Monomer recycle process for fluid phase in-line blending of polymers

ABSTRACT

A monomer recycle process for fluid phase in-line blending of polymers is provided. In one form, the monomer recycle process includes providing a first group (G1) of one or more reactor trains and a second group (G2) of one or more reactor trains and one or more separators fluidly connected to G1 and one separator fluidly connected to G2; polymerizing in each reactor train of G1 and G2 olefin monomers to form homogenous fluid phase polymer-monomer mixtures wherein each of the G1 and G2 reactor trains have at least one common monomer; passing the reactor effluents from the one or more G1 reactor trains through the one or more G1 separators to separate a monomer-rich phase from a polymer-enriched phase; passing the polymer-enriched phase and the reactor effluents from the one or more G2 reactor trains into the G2 separator (separator-blender) to separate another monomer-rich phase from a polymer-rich blend; recycling to one or more G1 reactor trains the separated monomer-rich phase from the one or more G1 separators; and recycling to one or more G2 reactor trains the separated monomer-rich phase from the G2 separator. The polymer-rich blend is conveyed to a downstream finishing stage for further monomer stripping, drying and/or pelletizing to form a polymer product blend.

CROSS-REFERENCE TO RELATED APPLICATIONS

This application that claims priority to U.S. Provisional ApplicationNo. 60/905,247 filed on Mar. 6, 2007, herein incorporated by referencein its entirety.

FIELD

The present disclosure relates to the field of polymer blending. It moreparticularly relates to a monomer recycle process for in-line blendingof polyolefin-based polymers in the fluid phase. Still moreparticularly, the present disclosure relates to a simplified monomerrecycle process for polymerization plants making in-line polymer blendsof homopolymers and copolymers in the fluid phase.

BACKGROUND

Polymer blends may be made by a variety of methods. A flexible butexpensive off-line process of making polymer blends uses solid polymersas starting materials, typically outside the polymerization process thatproduced the polymer blend components. The polymer blend components aretypically first melted or dissolved in a solvent and then blended. Theseprocesses are known as melt-blending and off-line solution blending,respectively. In melt blending, the solid, often pelletized or baled,polymer blend components are first melted and then blended together intheir molten state. One of the difficulties presented by melt blendingis the high viscosity of molten polymers, which makes blending of two ormore polymers difficult and often imperfect on the molecular level. Insolution off-line blending, the solid, often pelletized, polymer blendcomponents are first dissolved in a suitable solvent to form a polymersolution, and then two or more polymer solutions are blended together.After blending, solution blending requires the extraction of solventfrom the blend and drying of the blended polymer. Solution blending canovercome the viscosity issue associated with melt blending, but isexpensive due to the need for redissolving the polymer blend componentsand due to the cost of solvent handling.

The common feature of both melt blending and off-line solution blendingis that the polymer blending components are made in separate plants andthe solid polymers then are reprocessed either in a molten or in adissolved state to prepare the final polymer blend. In fact, theseoff-line blending processes are often operated by so-called compounders,generally independent of the manufacturers of the polymer blendcomponents. These processes add considerable cost to the cost of thefinal polymer blend. The production and full polymer recovery inseparate plants and subsequent reprocessing increases the costs ofproducing such blends because of the need for duplicate polymer recoverylines and because of the need for separate blending facilities and theenergy associated with their operations. Off-line solution blending alsorequires extra solvent, and facilities for polymer dissolution andsolvent recovery-recycle. Substantial reprocessing costs could be savedif the polymer blends could be made in one integrated polymerizationplant in-line, i.e. before the recovery and pelletizing of the solidpolymer blend components.

The disadvantage of a separate polyolefin blending plant associated withthe melt blending and off-line solution blending processes is alleviatedwith the prior art method of in-line solution blending of polymers usinga series reactor configuration. Utilizing the series reactorconfiguration, product blending may be accomplished in the solutionpolymerization reactor itself when the effluent of the first solutionpolymerization reactor is fed into the second reactor operating atdifferent conditions with optionally different catalyst and monomer feedcomposition. Referring to the two-stage series reactor configuration ofFIG. 1 (prior art), the two different polymers made in the first andsecond reactor stages are blended in the second stage yielding a blendedpolymer product leaving the second reactor. Such reactor seriesconfiguration may be further expanded into more than a two-stage seriesconfiguration (three or more reactors in series). Generally, a series ofn reactors may produce a blend with as many as n components or even morepresent in the effluent of the last reactor. Note that in principle,more than n components may be produced and blended in n reactors by, forexample, using more than one catalyst or by utilizing multiple zonesoperating at different conditions in one or more reactors of the seriesreactor cascade. While mixing in the downstream reactor(s) provides goodproduct mixing, particularly when the reactors are equipped with mixingdevices, e.g., mechanical stirrers, such series reactor configurationand operation presents a number of practical process and product qualitycontrol problems due to the close coupling of the reactors in thecascade. One of the most important difficulties in commercial practiceis ensuring proper blend and monomer ratios to deliver consistent blendquality. Additional complications arise when the blend components havedifferent monomer compositions, particularly when they have differentmonomer pools, such as in the case of blending different copolymers orin the case of blending homo- and copolymers. Since the monomer streamsare blended, there is no option for their separate recovery and recyclemandating costly monomer separations in the monomer recycle lines.

The above-outlined issues with series reactor operations are apparent tothose skilled in the art of chemical engineering. These difficulties areparticularly significant in polymerization because unlike insmall-molecule syntheses, reactor conditions determine not only reactorproductivities related to product blend ratio, but also productproperties related to controlling the quality of the polymer blendcomponents. For example, FIGS. 2 and 3 show how reactor temperature andpressure affect polymer properties of fundamental importance, such asmolecular weight (MW) and melting behavior. Surprisingly, we found thatmonomer conversion in the reactor also influences these critical productattributes (see FIG. 4). Since in a series reactor cascade the effluentof an upstream reactor flows into the next downstream member of thereactor cascade, the residence time, catalyst concentration, and monomercomposition and thus monomer conversion in the downstream reactor cannotbe adjusted independently of the operating conditions (particularly ofthe flow rate) of the upstream reactor. Because of this close andinherent coupling of operating regimes in the reactors of the seriescascade, the correlations depicted in FIGS. 2, 3, and 4 further reducethe controllability, flexibility, and thus the usefulness of the in-lineblending method in a series reactor configuration. Ultimately, thisgreatly reduces the number of blend products that can be made in such aseries reactor cascade and makes the blend quality difficult to control.

Applying parallel reactors can overcome the disadvantages related to thedirect coupling of the polymerization reactors in an in-line polymerblending applying series reactors. While production flexibility isincreased, a parallel reactor arrangement necessitates the installationof blending vessels increasing the cost of the process. As disclosed inU.S. Patent Application No. 60/876,193 filed on Dec. 20, 2006, hereinincorporated by reference in its entirety, an improved in-line processfor blending polymers has been developed to improve blend quality andreduce the capital and operating costs associated with a combinedpolymerization and blend plant.

One problem associated with the fluid phase in-line blending processutilizing two or more parallel reactor trains and the two or morehigh-pressure separators fluidly connected to the two or more reactortrains configured in parallel is that the polymerization processes donot convert 100% of the feed monomers in a single pass through theparallel reactor trains. Because the monomers are of high value, theunconverted monomer streams are typically recovered by separating themfrom the polymeric products and recycled back to the polymerizationreactor bank. However, when the monomer-rich phase from a high-pressureseparator contains two or more different monomers, separation andrecycle become more challenging and complex. In particular, when themonomer-rich phase from a high pressure separator contains a mixture ofa monomer stream from a homopolymerization parallel reactor train and acomonomer stream from a copolymerization parallel reactor trains,recycle of the monomer and comonomer streams becomes particularlychallenging due to the fact that the monomer composition of the recyclestream is often too far from the desired reactor feed composition. Thisimbalance can be best quantified by the excess monomer component flowpresent in the recycle monomer stream. In order to allow the rebalancingof the monomer feed composition, the monomers present in the combinedrecycle stream are typically split into individual pure monomer feeds.These individual monomer recycle streams then are combined with the makeup monomer streams at a controlled rate to provide for the reactor feedsat the required monomer feed rate and composition. Such monomerseparation trains, however, are capital intensive and expensive tooperate. They are particularly expensive for light olefinic monomerrecycle streams comprising aliphatic olefins of two to four carbon atomsdue to their low boiling points and thus due to the need for cryogenicdistillation. Hence, there is a need for new inventive processes thatminimize or even eliminate such expensive monomer separations in themonomer recycle train of in-line blending processes.

There are a variety of possible options for recovering and recycling themonomers. One option, outside the scope of the present disclosure,recovers the monomers in separate trains that do not contact each otherand combine the monomer and comonomer streams before, during, or afterthe monomer separation from the polymers. The monomer and comonomerstreams are kept separate and may be recycled to their correspondingpolymerization reactor trains without concerns about contaminating thefeed of the homopolymerization train by the comonomer of thecopolymerization train. While this option eliminates the need for andthe cost associated with separation and recovery of the individualmonomers, it requires parallel feed-product separation trains for eachreactor train. It also handles and blends the undiluted, highly viscouspolymer blend components in their molten state, as opposed to theirdiluted fluid state. Both of these requirements increase cost anddecrease blend quality.

FIG. 5 depicts the process flow diagram of another prior art monomerrecycle process used in conjunction with fluid phase in-line blending ofpolymers. Referring to FIG. 5, the reactor effluents (1) and (2) fromthe two parallel reactor trains (1) and (2) are blended in the singleseparator-blender that also separates a monomer-rich phase from apolymer-rich phase. The monomer-rich phase emerging from the top of theseparator-blender comprises a mixture of the monomers of reactor train(1) and the monomers from reactor train (2). This combined monomerstream is then separated into its individual monomer components beforethose components are recycled to the proper reactor trains of thein-line polymer blending process. This prior art monomer separation andrecycle method provides ultimate flexibility in the recycle train sinceeach monomer component is recovered individually and can be directed atthe desired rate to the appropriate reactor train(s). However, it alsoapplies a complex and expensive monomer separation train. A recycleprocess that can operate without the full recovery of each individualmonomer components would be cost advantaged, simpler to operate andmaintain, and thus desirable.

Besides separating the monomer components present in the monomer recyclestream emerging from the separator-blender, heavier impurities, such assolvents used in the catalyst feed, oligomeric side products, excesscatalyst killer(s), etc. need to be separated and purged from themonomer recycle stream in order to prevent their buildup in the reactortrains. Again, referring to FIG. 5, separation tower (1) separates thecombined monomer-comonomer stream from solvents, oligomers and otherheavies, which emerge from the bottom of tower (1). Separation tower (2)further separates the monomer-containing stream emerging from the top oftower (1). Monomer (1) emerges from the bottom and monomer (2) emergesfrom the top of tower (2). The monomer (1) bottom stream from separationtower (2) is recycled back to polymerization parallel reactor train (1)and the monomer (2) top stream from separation tower (2) is recycledback to polymerization parallel reactor train (2). It should beunderstood that in the case of monomer recycle streams comprising morethan two monomers, more than two separation towers may be necessary torecover the individual monomer components. These chilled separationtowers and associated hardware are expensive to install from a capitalequipment standpoint and expensive to operate from utility standpoint.The separation towers also increase maintenance costs associated withthe polymerization-blending plant. In addition, the separation processstep also increases the overall in-line blending process complexity andthus the probability of process upsets. Therefore, there is a need foran improved method that can reduce the complexity and cost of monomerrecycle related to the in-line polymer blending process.

SUMMARY

Provided is a monomer recycling process for use with the fluid phasein-line blending of polymers.

One aspect of the present disclosure relates to an advantageous monomerrecycle process for use with fluid phase in-line blending of polymerscomprising: (a) providing two or more reactor trains configured inparallel and two or more high-pressure separators downstream fluidlyconnected to the two or more reactor trains configured in parallel; (b)contacting in the two or more reactor trains configured in parallel 1)olefin monomers having two or more carbon atoms, 2) one or more catalystsystems, 3) optional one or more comonomers, 4) optional one or morescavengers, and 5) optional one or more diluents or solvents, wherein atleast one of the reactor trains configured in parallel is at atemperature above the solid-fluid phase-transition temperature of thepolymerization system and a pressure no lower than 10 MPa below thecloud point pressure of the polymerization system and less than 1500MPa, wherein the polymerization system for each reactor train is in itsdense fluid state and comprises the olefin monomers, any comonomerpresent, any diluent or solvent present, any scavenger present, and thepolymer product, and wherein the catalyst system for each reactor traincomprises one or more catalyst precursors, one or more activators, andoptionally, one or more catalyst supports; (c) forming an unreducedreactor effluent including a homogenous fluid phase polymer-monomermixture in each parallel reactor train, wherein one or more of theparallel reactor trains define a first group (G1) of reactor trains andanother one or more of the parallel reactor trains define a second group(G2) of reactor trains, wherein the number of monomers (N) in the feedmonomer pools for G1 (N(G1)), G2 (N(G2)), and for the combined feedmonomer pool of G1 and G2 (N(G1+G2)) are related as follows:N(G1+G2)=N(G2) and N(G1)≦N(G2), and wherein the feed monomer pool ineach reactor train of G2 are the same; (d) passing the unreduced reactoreffluents from one or more of the G1 reactor trains through one or morehigh-pressure separators, maintaining the temperature and pressurewithin the one or more high-pressure separators above the solid-fluidphase transition point but below the cloud point pressure andtemperature to form fluid-fluid two-phase systems in each of the one ormore high-pressure separators with each two-phase system comprising apolymer-enriched phase and a monomer-rich phase, and separating themonomer-rich phase from the polymer-enriched phase in each of the one ormore high-pressure separators to form one or more separated monomer-richphases and one or more polymer-enriched phases; (e) passing the one ormore polymer-enriched phases from the one or more high-pressureseparators of (d), any unreduced reactor effluents of G1 not passingthrough the one or more high-pressure separators of G1, and theunreduced reactor effluents of the G2 reactor trains through anotherhigh-pressure separator for product blending and product-feedseparation; (f) maintaining the temperature and pressure within theanother high-pressure separator of (e) above the solid-fluid phasetransition point but below the cloud point pressure and temperature toform a fluid-fluid two-phase system comprising a polymer-rich blendphase and a monomer-rich phase; (g) separating the monomer-rich phasefrom the polymer-rich blend phase in the another high-pressure separatorof (e) to form a polymer-rich blend and a separated monomer-rich phase;(h) recycling the one or more separated monomer-rich phases of (d) toone or more reactor trains of G1; and (i) recycling the separatedmonomer-rich phase of (g) to one or more reactor trains of G2.

Another aspect of the present disclosure relates to an advantageousmonomer recycle process for use with fluid phase in-line blending ofpolymers comprising: (a) providing two or more reactor trains configuredin parallel and a high-pressure separator downstream fluidly connectedto the two or more reactor trains configured in parallel; (b) contactingin the two or more reactor trains configured in parallel 1) olefinmonomers having two or more carbon atoms, 2) one or more catalystsystems, 3) optional one or more comonomers, 4) optional one or morescavengers, and 5) optional one or more diluents or solvents, wherein atleast one of the reactor trains configured in parallel is at atemperature above the solid-fluid phase-transition temperature of thepolymerization system and a pressure no lower than 10 MPa below thecloud point pressure of the polymerization system and less than 1500MPa, wherein the polymerization system for each reactor train is in itsdense fluid state and comprises the olefin monomers, any comonomerpresent, any diluent or solvent present, any scavenger present, and thepolymer product, and wherein the catalyst system for each reactor traincomprises one or more catalyst precursors, one or more activators, andoptionally, one or more catalyst supports; (c) forming a reactoreffluent including a homogeneous fluid phase polymer-monomer mixture ineach parallel reactor train, wherein one or more of the parallel reactortrains define a first group (G1) of reactor trains and another one ormore of the parallel reactor trains define a second group (G2) ofreactor trains, wherein the number of monomers (N) in the feed monomerpools for G1 (N(G1)), G2 (N(G2)), and for the combined feed monomer poolof G1 and G2 (N(G1+G2)) are related as follows: N(G1+G2)=N(G2) andN(G1)≦N(G2), and wherein the monomer pool in each reactor train of G2are the same; (d) passing the reactor effluent comprising thehomogeneous fluid phase polymer-monomer mixture from each reactor trainof G1 and G2 through the high-pressure separator for product blendingand product-feed separation; (e) maintaining the temperature andpressure within the high-pressure separator above the solid-fluid phasetransition point but below the cloud point pressure and temperature toform a fluid-fluid two-phase system comprising a polymer-rich blendphase and a monomer-rich phase; (f) separating the monomer-rich phasefrom the polymer-rich blend phase in the high-pressure separator to forma polymer-rich blend and a separated monomer-rich phase; and (g)recycling the separated monomer-rich phase of (f) to one or more reactortrains of G2.

Still another aspect of the present disclosure relates to anadvantageous monomer recycle process for use with fluid phase in-lineblending of polymers comprising: (a) providing a first group (G1) of oneor more parallel reactor trains and a second group (G2) of one or moreparallel reactor trains, wherein G1 polymerizes a polypropylenehomopolymer and G2 polymerizes a polypropylene copolymer, and whereinone high-pressure separator is fluidly connected to G1 and anotherhigh-pressure separator is fluidly connected to G2; (b) contacting ineach of the parallel reactor trains of G1 and G2: 1) propylene monomer,2) one or more catalyst systems, 3) one or more comonomers in each ofthe parallel reactor trains of G2, 4) optional one or more scavengers,and 5) optional one or more diluents or solvents, wherein each of theparallel reactor trains for G1 and G2 are at a temperature above thesolid-fluid phase-transition temperature of the polymerization systemand a pressure no lower than 10 MPa below the cloud point pressure ofthe polymerization system and less than 1500 MPa and have at least onecommon monomer, wherein the polymerization system for each parallelreactor train of G1 and G2 is in its dense fluid state and comprisespropylene monomer, any diluent or solvent present, any scavengerpresent, and the polymer product, and the polymerization system for thesecond reactor train further comprises one or more comonomers, andwherein the catalyst system for each parallel reactor train of G1 and G2comprises one or more catalyst precursors, one or more activators, andoptionally, one or more catalyst supports; (c) forming an unreducedreactor effluent including a homogenous fluid phase polymer-monomermixture in each parallel reactor train of G1 and G2; (d) passing theunreduced reactor effluents from the one or more of the G1 parallelreactor trains through a high-pressure separator while maintaining thetemperature and pressure within the high-pressure separator above thesolid-fluid phase transition point but below the cloud point pressureand temperature to form a fluid-fluid two-phase systems comprising apolymer-enriched phase and a monomer-rich phase, and separating themonomer-rich phase from the polymer-enriched phase in the high-pressureseparator to form a separated monomer-rich phase and a polymer-enrichedphase; (e) passing the polymer-enriched phase from the high-pressureseparator of (d) and the unreduced reactor effluents of the G2 reactortrains through another high-pressure separator for product blending andproduct-feed separation; (f) maintaining the temperature and pressurewithin the another high-pressure separator of (e) above the solid-fluidphase transition point but below the cloud point pressure andtemperature to form a fluid-fluid two-phase system comprising apolymer-rich blend phase and a monomer-rich phase; (g) separating themonomer-rich phase from the polymer-rich blend phase in the anotherhigh-pressure separator of (e) to form a polymer-rich blend and aseparated monomer-rich phase; (h) recycling the separated monomer-richphase of (d) to the one or more reactor trains of G1; and (i) recyclingthe separated monomer-rich phase of (g) to one or more reactor trains ofG2.

Still yet another aspect of the present disclosure relates to anadvantageous monomer recycle process for use with fluid phase in-lineblending of polymers comprising: (a) providing a first group (G1) of oneor more reactor trains and a second group (G2) of one or more reactortrains, wherein G1 polymerizes a polypropylene homopolymer and G2polymerizes a polypropylene copolymer, wherein a high-pressure separatoris fluidly connected to G1 and G2, and wherein the G1 reactor trains areconfigured parallel to and fluidly connected to the G2 reactor trains;(b) contacting in each of the reactor trains of G1 and G2: 1) propylenemonomer, 2) one or more catalyst systems, 3) one or more comonomers ineach of the reactor trains of G2, 4) optional one or more scavengers,and 5) optional one or more diluents or solvents, wherein each of thereactor trains for G1 and G2 are at a temperature above the solid-fluidphase-transition temperature of the polymerization system and a pressureno lower than 10 MPa below the cloud point pressure of thepolymerization system and less than 1500 MPa and have at least onecommon monomer, wherein the polymerization system for each reactor trainof G1 and G2 is in its dense fluid state and comprises propylenemonomer, any diluent or solvent present, any scavenger present, and thepolymer product, and the polymerization system for the second reactortrain further comprises one or more comonomers, and wherein the catalystsystem for each reactor train of G1 and G2 comprises one or morecatalyst precursors, one or more activators, and optionally, one or morecatalyst supports; (c) forming a reactor effluent including ahomogeneous fluid phase polymer-monomer mixture in each reactor train ofG1 and G2; (d) passing the reactor effluents from each reactor train ofG1 and G2 through the high-pressure separator for product blending andproduct-feed separation; (e) maintaining the temperature and pressurewithin the high-pressure separator above the solid-fluid phasetransition point but below the cloud point pressure and temperature toform a fluid-fluid two-phase system comprising a polymer-rich blendphase and a monomer-rich phase; (f) separating the monomer-rich phasefrom the polymer-rich blend phase in the high-pressure separator to forma polymer-rich blend and a separated monomer-rich phase; and (g)recycling the separated monomer-rich phase of (f) to the one or morereactor trains of G2.

These and other features and attributes of the disclosed monomer recycleprocesses for fluid phase in-line blending of polymers and theiradvantageous applications and/or uses will be apparent from the detaileddescription that follows, particularly when read in conjunction with thefigures appended hereto.

BRIEF DESCRIPTION OF THE DRAWINGS

To assist those of ordinary skill in the relevant art in making andusing the subject matter hereof, reference is made to the appendeddrawings, wherein:

FIG. 1 presents the process for the production of polymer blends in atwo-stage series reactor configuration (prior art);

FIG. 2 presents the effect of polymerization temperature on themolecular weight and melting point of polypropylene made insupercritical polypropylene using MAO-activated(μ-dimethylsilyl)bis(2-methyl-4-phenylindenyl)zirconium dichloride(Q-Zr-MAO) catalyst at 207 MPa (30 kpsi);

FIG. 3 presents the effect of polymerization pressure on the molecularweight and melting point of polypropylene made in supercriticalpropylene using MAO-activated(μ-dimethylsilyl)bis(2-methyl-4-phenylindenyl)zirconium dichloride(Q-Zr-MAO) catalyst at 120 and 130° C.;

FIG. 4 presents the effect of propylene conversion in the polymerizationof supercritical propylene using MAO-activated(μ-dimethylsilyl)bis(2-methyl-4-phenylindenyl)zirconium dichloride(Q-Zr-MAO) at 130° C. and 69 and 138 MPa (10 or 20 kpsi, respectively);

FIG. 5 presents a prior art monomer recycle process with full recoveryof individual monomer components used in conjunction with a fluid phasein-line polymer blending process;

FIG. 6 presents an exemplary monomer recycle process for fluid phasein-line blending of polymers with one separator vessel(separator-blender), the combined monomer recycle stream of reactortrains (1) and (2) recycled to train (2) without recovering individualmonomer components;

FIG. 7 presents an exemplary monomer recycle process for fluid phasein-line blending of polymers with one separator vessel(separator-blender) and with additive/polymer blending component option,the combined monomer recycle stream of reactor trains (1) and (2)recycled to train (2) without recovering individual monomer components;

FIG. 8 presents an exemplary monomer recycle process for fluid phasein-line blending of polymers with two separation vessels;

FIG. 9 presents another exemplary monomer recycle process for fluidphase in-line blending of polymers with two separation vessels with aproduct effluent buffer tank that also serves as monomer/productseparator for improved blend ratio control;

FIG. 10 presents an exemplary monomer recycle process for fluid phasein-line blending of polymers with two separation vessels and withadditive/polymer blending component option;

FIG. 11 presents an exemplary monomer recycle process for fluid phasein-line blending of polymers with two separation vessels, one alsoserving as buffer, and with additive/polymer blending component option;

FIG. 12 presents cloud point isotherms for Polymer Achieve™ 1635;

FIG. 13 presents cloud point isotherms for Polymer PP 45379 dissolved inbulk propylene;

FIG. 14 presents cloud point isotherms for Polymer PP 4062 dissolved inbulk propylene;

FIG. 15 presents cloud point isotherms for Polymer Achieve™ 1635dissolved in bulk propylene;

FIG. 16 presents cloud point isotherms for Polymer PP 45379 dissolved inbulk propylene;

FIG. 17 presents cloud point isotherms for Polymer PP 4062 dissolved inbulk propylene;

FIG. 18 presents a comparison of isopleths for PP 45379, Achieve™ 1635,and PP 4062 dissolved in bulk propylene;

FIG. 19 presents a comparison of isopleths for Achieve™ 1635 andliterature data as described in J. Vladimir Oliveira, C. Dariva and J.C. Pinto, Ind. Eng, Chem. Res. 29, 2000, 4627;

FIG. 20 presents a comparison of isopleths for PP 45379 and literaturedata as described in J. Vladimir Oliveira, C. Dariva and J. C. Pinto,Ind. Eng, Chem. Res. 29 (2000), 4627;

FIG. 21 presents a comparison of isopleths for PP 4062 and literaturedata as described in J. Vladimir Oliveira, C. Dariva and J. C. Pinto,Ind. Eng, Chem. Res. 29, 2000, 4627;

FIG. 22 presents a basic phase diagram for mixture of propylene monomerwith selected polymers (isotactic polypropylene—iPP, syndiotacticpolypropylene—sPP, atactic polypropylene—aPP, or propylene-ethylenecopolymer);

FIG. 23 presents a comparison of the density of supercritical propyleneat 137.7° C. with liquid propylene at 54.4° C.;

FIG. 24 presents an operating regime in accordance with the processdisclosed herein for a reactor operating in a single liquid phase;

FIG. 25 presents an operating regime in accordance with the processdisclosed herein for a reactor operating in a liquid-liquid phase; and

FIG. 26 presents an operating regime in accordance with the processdisclosed herein for a gravity separator.

DEFINITIONS

For the purposes of this invention and the claims thereto:

A catalyst system is defined to be the combination of one or morecatalyst precursor compounds and one or more activators. Althoughthemselves are not catalytically active (need to be combined with anactivator to become active), the catalyst precursor compounds are oftenreferred to as catalysts in the art of polymerization. Any part of thecatalyst system can be optionally supported on solid particles, in whichcase the support is also part of the catalyst system.

Dense fluids are defined as fluid media in their liquid or supercriticalstate with densities greater than 300 kg/m³. Note that gas-phase fluidsare excluded from the group of dense fluids.

Solid-fluid phase transition temperature is defined as the temperatureat which a solid polymer phase separates from the polymer-containingdense fluid medium at a given pressure. The solid-fluid phase transitiontemperature can be determined by temperature reduction starting fromtemperatures at which the polymer is fully dissolved in the dense fluidreaction medium. At the onset of the formation of a solid polymer phase,the homogeneous fluid medium becomes turbid, which can be observed byeye or can be detected by shining a laser through the medium anddetecting the sudden increase of light scattering as described in J.Vladimir Oliveira, C. Dariva and J. C. Pinto, Ind. Eng, Chem. Res. 29(2000) 4627.

Solid-fluid phase transition pressure is defined as the pressure atwhich a solid polymer phase separates from the polymer-containing fluidmedium at a given temperature. The solid-fluid phase transition pressurecan be determined by pressure reduction at constant temperature startingfrom pressures at which the polymer is fully dissolved in the fluidreaction medium. At the onset of the formation of a solid polymer phase,the homogeneous fluid medium becomes turbid, which can be observed byeye or can be detected by shining a laser through the medium anddetecting the sudden increase of light scattering as described in J.Vladimir Oliveira, C. Dariva and J. C. Pinto, Ind. Eng, Chem. Res. 29(2000) 4627.

The cloud point is defined as the pressure below which, at a giventemperature, the polymer-containing homogeneous fluid medium becomesturbid upon pressure reduction at constant temperature as described inJ. Vladimir Oliveira, C. Dariva and J. C. Pinto, Ind. Eng, Chem. Res. 29(2000) 4627. For purposes of this invention and the claims thereto, thecloud point is measured by shining a helium laser through the selectedpolymerization system in a cloud point cell onto a photocell andrecording the pressure at the onset of rapid increase in lightscattering for a given temperature.

A higher α-olefin is defined as an α-olefin having four or more carbonatoms.

Polymerization encompasses any polymerization reaction such ashomopolymerization and copolymerization.

Copolymerization encompasses any polymerization reaction of two or moremonomers.

The new numbering scheme for the Periodic Table Groups is used aspublished in CHEMICAL AND ENGINEERING NEWS, 63(5), 27 (1985).

When a polymer is referred to as comprising an olefin, the olefinpresent in the polymer is the polymerized form of the olefin.

An oligomer is defined to be compositions having 2-75 monomer units.

A polymer is defined to be compositions having 76 or more monomer units.

A series reactor cascade (also referred to as series reactorconfiguration or reactors in series) includes two or more reactorsconnected in series, in which the effluent of an upstream reactor is fedto the next reactor downstream in the reactor cascade. Besides theeffluent of the upstream reactor(s), the feed of any reactor can beaugmented with any combination of additional monomer, catalyst,scavenger, or solvent fresh or recycled feed streams.

Reactor train or reactor branch or reactor leg refers to a singlepolymerization reactor or to a group of polymerization reactors of thein-line blending process disclosed herein that produces a single polymerblend component. If the reactor train contains more than one reactor,the reactors are arranged in a series configuration within the train.The need for having more than one reactor in a reactor train may, forexample, arise when an in-line blend component cannot be produced at thedesired rate economically in a single reactor but there could be alsoreasons related to blend component quality, such as molecular weight orcomposition distribution, etc. Since a reactor train can comprisemultiple reactors and/or reactor zones in series, the single blendcomponent produced in a reactor train may itself be a polymer blend ofpolymeric components with varying molecular weights and/or compositions.However, in order to simplify the description of different embodimentsof the in-line blending processes disclosed herein, the polymericproduct of a reactor train is referred to simply as blend component orpolymeric blend component regardless of its molecular weight and/orcompositional dispersion. For the purpose of defining the process of thepresent invention, parallel reactors will be always considered asseparate reactor trains even if they produce essentially the samein-line blend component. Also, spatially separated, parallel reactionzones that do not exchange or mix reaction mixtures by, for example,pump-around loops, or by other recirculation methods, will be consideredas separate parallel reactor trains even when those parallel zones arepresent in a common shell and fall within the in-line blending processdisclosed herein.

Reactor bank refers to the combination of all polymerization reactors inthe polymerization section of the in-line polymer blending process isdisclosed herein. A reactor bank may comprise one or more reactortrains.

A parallel reactor configuration includes two or more reactors orreactor trains connected (also referred to as fluidly connected) inparallel. A reactor train, branch, or leg may include one reactor oralternatively more than one reactor configured in a seriesconfiguration. For example, a reactor train may include two, or three,or four, or more reactors in series. The entire parallel reactorconfiguration of the polymerization process disclosed herein, i.e., thecombination of all parallel polymerization reactor trains forms thereactor bank.

Monomer recycle ratio refers to the ratio of the amount of recycledmonomer fed to the reactor divided by the total (fresh plus recycled)amount of monomer fed to the reactor.

Polymerization system is defined to be the monomer(s) plus comonomer(s)plus polymer(s) plus optional inert solvent(s)/diluent(s) plus optionalscavenger(s). Note that for the sake of convenience and clarity, thecatalyst system is always addressed separately in the present discussionfrom other components present in a polymerization reactor. In thisregard, the polymerization system is defined here narrower thancustomary in the art of polymerization that typically considers thecatalyst system as part of the polymerization system. In the currentdefinition, the mixture present in the polymerization reactor and in itseffluent is composed of the polymerization system plus the catalystsystem.

A homogeneous polymerization system contains all of its componentsdispersed and mixed on a molecular scale. In our discussions,homogeneous polymerization systems are meant to be in their dense fluid(liquid or supercritical) state. Note that our definition of thepolymerization system does not include the catalyst system, thus thecatalyst system may or may not be homogeneously dissolved in thepolymerization system. A homogeneous system may have regions withconcentration gradients, but there would be no sudden, discontinuouschanges of composition on a micrometer scale within the system. Inpractical terms, a homogeneous polymerization system has all of itscomponents in a single dense fluid phase. Apparently, a polymerizationsystem is not homogeneous when it is partitioned to more than one fluidphase or to a fluid and a solid phase. The homogeneous fluid state ofthe polymerization system is represented by the single fluid region inits phase diagram.

Pure substances, including all types of hydrocarbons, can exist ineither a subcritical, or supercritical state, depending on theirtemperature and pressure. To be in the supercritical state, a substancemust have a temperature above its critical temperature (Tc) and apressure above its critical pressure (Pc). Mixtures of hydrocarbons,including mixtures of monomers, polymers, and optionally inert solvents,have pseudo-critical temperatures (Tc) and pseudo-critical pressures(Pc), which for many systems can be approximated bymole-fraction-weighted averages of the corresponding critical properties(Tc or Pc) of the mixture's components. Mixtures with a temperatureabove their pseudo-critical temperature and a pressure above theirpseudo-critical pressure will be said to be in a supercritical state orphase, and the thermodynamic behavior of supercritical mixtures will beanalogous to supercritical pure substances. For purposes of thisinvention, the critical temperatures (Tc) and critical pressures (Pc) ofcertain pure substances relevant to the current invention are those thatfound in the Handbook of Chemistry and Physics, David R. Lide,Editor-in-Chief, 82nd edition 2001-2002, CRC Press, LLC. New York, 2001.In particular, the Tc and Pc of selected substances are:

Name Tc (K) Pc (MPa) Hexane 507.6 3.025 Isobutane 407.8 3.64 Ethane305.3 4.872 Cyclobutane 460.0 4.98 Cyclopentane 511.7 4.51 1-Butene419.5 4.02 1-pentene 464.8 3.56 Pentane 469.7 3.37 Benzene 562.05 4.8951-hexene 504.0 3.21 Propane 369.8 4.248 Toluene 591.8 4.11 Methane190.56 4.599 Butane 425.12 3.796 Ethylene 282.34 5.041 Propylene 364.94.6 Cyclopentene 506.5 4.8 Isopentane 460.4 3.38 Cyclohexane 553.8 4.08Heptane 540.2 2.74 273.2 K = 0° C.

The following abbreviations are used: Me is methyl, Ph is phenyl, Et isethyl, Pr is propyl, iPr is isopropyl, n-Pr is normal propyl, Bu isbutyl, iBu is isobutyl, tBu is tertiary butyl, p-tBu is para-tertiarybutyl, TMS is trimethylsilyl, TIBA is tri-isobutylaluminum, MAO ismethylaluminoxane, pMe is para-methyl, flu is fluorenyl, cp iscyclopentadienyl.

By continuous is meant a system that operates (or is intended tooperate) without interruption or cessation. For example, a continuousprocess to produce a polymer would be one where the reactants arecontinually introduced into one or more reactors and polymer product iscontinually withdrawn.

Slurry polymerization refers to a polymerization process in whichparticulate, solid polymer (e.g., granular) forms in a dense fluid or ina liquid/vapor polymerization medium. The dense fluid polymerizationmedium can form a single or two fluid phases, such as liquid,supercritical fluid, or liquid/liquid, or supercriticalfluid/supercritical fluid, polymerization medium. In a liquid/vaporpolymerization medium, the polymer resides in the liquid (dense) phase.Slurry polymerization processes typically employ heterogeneous catalystparticles, such as Ziegler-Natta catalysts or supported metallocenecatalysts, and the like. The solid polymeric product typically adheresto the heterogeneous solid catalyst particles thus forming a slurryphase. Slurry polymerization processes operate below the solid-fluidphase transition temperature of the polymerization system.

Solution polymerization refers to a polymerization process in which thepolymer is dissolved in a liquid polymerization system, such as an inertsolvent or monomer(s) or their blends. Solution polymerization comprisesa homogeneous liquid polymerization system in the reactor. Thetemperature of a liquid polymerization system is below of itssupercritical or pseudo supercritical temperature, thus solutionpolymerizations are performed below the supercritical temperature and/orpressure.

Supercritical polymerization refers to a polymerization process in whichthe polymerization system is in its dense supercritical or pseudosupercritical state, i.e. when the density of the polymerization systemis above 300 g/L and its temperature and pressure are above thecorresponding critical values.

Bulk polymerization refers to a polymerization process in which thedense fluid polymerization system contains less than 40 wt %, or lessthan 30 wt %, or less than 20 wt %, or less than 10 wt %, or less than 5wt %, or less than 1 wt % of inert solvent. Inert solvents arecharacterized by their lack of incorporation into the product polymerchain. In the production of polyolefins, solvents are typicallyhydrocarbons comprising 4 to 20 carbon atoms, advantageously 5 to 10, or5 to 8 carbon atoms. Note that the polymerization system may alsocontain inert diluents that do not incorporate into the product polymerchain. They are typically introduced as impurities present in themonomer feeds. For the purpose of the current disclosure, the inertdiluents are considered separately from the inert solvents, the latterof which are added intentionally for their ability to keep the polymericproducts in their dissolved state.

A homogeneous polymerization system contains all of its componentsdispersed and mixed on a molecular scale. In our discussions,homogeneous polymerization systems are meant to be in their dense fluid(liquid or supercritical) state. Note that our definition of thepolymerization system does not include the catalyst system, thus thecatalyst system may or may not be homogeneously dissolved in thepolymerization system. A homogeneous system may have regions withconcentration gradients, but there would be no sudden, discontinuouschanges of composition on a micrometer scale within the system as it isthe case when, for example, solid polymer-containing particles aresuspended in a dense fluid. In practical terms, a homogeneouspolymerization system has all of its components in a single dense fluidphase. Apparently, a polymerization system is not homogeneous when it ispartitioned to more than one fluid phase or to a fluid and a solidphase. The homogeneous fluid state of the polymerization system isrepresented by the single fluid (liquid or supercritical fluid) regionin its phase diagram.

Homogeneous supercritical polymerization refers to a polymerizationprocess in which the polymer is dissolved in a dense supercritical fluidpolymerization medium, such as an inert solvent or monomer or theirblends in their supercritical state. As described above, when thesupercritical fluid polymerization system contains less than 40 wt %, orless than 30 wt %, or less than 20 wt %, or less than 10 wt %, or lessthan 5 wt %, or less than 1 wt % of inert solvent and the polymer isdissolved in the dense supercritical fluid, the process is referred toas a bulk homogeneous supercritical polymerization process. Homogeneoussupercritical polymerization should be distinguished from heterogeneoussupercritical polymerizations, such as for example, supercritical slurryprocesses, the latter of which are performed in supercritical fluids butform solid polymer particulates in the polymerization reactor.Similarly, bulk homogeneous supercritical polymerization should bedistinguished from bulk solution polymerization, the latter of which isperformed in a liquid as opposed to in a supercritical polymerizationsystem.

An in-line blending process disclosed herein refers to one where thepolymerization and the polymer blending processes are integrated in asingle process and at least one of the polymerization trains operatesunder solution or homogeneous supercritical conditions. Although in-lineblending processes typically employ polymerization trains using solutionor homogeneous supercritical polymerization systems, one or more of thepolymerization trains may employ slurry polymerization systems,particularly bulk slurry polymerization systems. When the polymerizationbank includes one or more slurry polymerization trains, the effluents ofthose slurry trains are always heated above their solid-fluid transitionpoints and optionally pressurized before mixing them with the effluentsof other trains to enable fluid-phase mixing.

In-line polymer blend or in-line blend disclosed herein refers to amixture of two or more polymeric components, at least one of which isproduced under either homogeneous supercritical polymerizationconditions or homogeneous solution polymerization conditions. Thepolymeric components are produced internally in the in-line blendingprocess and are mixed in the same process without recovering them intheir solid state. Optionally, the in-line blends may also containadditives produced outside the invention process, such as plasticizers,UV stabilizers, antioxidants, etc., and off-line polymericadditives/modifiers in minor amounts, i.e., less than 50%, or less than40%, or less than 30%, or less than 20%, or less than 10%, or less than5%, or less than 1% by weight.

Single-pass conversion of monomer i in reactor train j is defined by thefollowing formula:

Single pass conversion of monomer i in reactor train j (%)=100×[(reactortrain j effluent flow rate in weight per hour)×(polymer productconcentration in the effluent of reactor train j in weightfraction)×(monomer i concentration in the polymer product made inreactor train j in weight fraction)]/[(reactor train j feed flow rate,including fresh and recycled, in weight per hour)×(monomer iconcentration in the feed of reactor train j, including fresh andrecycled, in weight fraction)].

Overall conversion of monomer i in reactor train j is defined by thefollowing formula:

Overall conversion of monomer i in reactor train j (%)=100×[(reactortrain j effluent flow rate in weight per hour)×(polymer productconcentration in the effluent of reactor train j in weightfraction)×(monomer i concentration in the polymer product made inreactor train j in weight fraction)]/[(reactor train j fresh feed flowrate, excluding recycled, in weight per hour)×(monomer i concentrationin the fresh feed of reactor train j, excluding recycled, in weightfraction)].

The total monomer feed rate of a given train is equal to the sum offresh (make up) monomer feed rate plus recycled monomer feed rate inweight per hour.

The flow (feed, effluent, recycle, purge etc. flow) rate of monomercomponent i in weight per hour is equal to the total flow (feed,effluent, recycle, purge etc. flow) rate in weight per hour multipliedby the weight fraction of monomer component i.

Conversion rate of monomer i in reactor train j expressed in weight perhour is equal to [(reactor train j effluent flow rate in weight perhour)×(polymer product concentration in the effluent of reactor train jin weight fraction)×(monomer i concentration in the polymer product madein reactor train j in a single pass in weight fraction)].

Monomer pool describes the number and nature of the monomers present notless than 1.0 wt %, or not less than 0.5 wt %, or not less than 0.1 wt%, or not less than 0.05 wt %, or not less than 0.01 wt %, or not lessthan 0.001 wt % in the process or in certain feeds or effluents orrecycle streams or vessels or products of the process. In general, onecan refer to the monomer pool of reactor r, or the monomer pool ofreactor train j, or the monomer pool in the feed or in the effluent ofreactor r or of reactor train j, or the monomer pool of the entireprocess, or the monomer pool of the product of a reactor train or of areactor group, or of the process, etc. While the monomer pool pertainsto the number and nature of the monomers it represents, it does notspeak to the quantity or concentration of the monomers of the polymerpool (except for setting a minimum/threshold level below which a monomeris not considered to be a part of the monomer pool). Two or more monomerpools containing the same and only the same monomers are considered tobe identical or the same even if the (higher than the minimum/threshold)concentration or the ratio of the monomers in the pools are not thesame. Naturally, a monomer is not present in a pool when itsconcentration is below a defined threshold and thus only considered tobe present in a monomer pool or considered to be part of a monomer poolwhen its concentration is higher than the defined threshold. Twoexamples of monomer pools are as follows. (1) The monomer pool of astream containing ethylene, propylene, and 1,5-hexadiene has threemembers: ethylene, propylene, and 1,5-hexadiene. (2) The combinedmonomer pools of reactor train 1 containing ethylene and propylene,reactor train 2 containing propylene and butene-1, and reactor train 3containing propylene has also three members: ethylene, propylene, andbutene-1. Note that the monomer pool of train 1 has two members:ethylene and propylene, the monomer pool of train 2 has also twomembers: propylene and butene-1, and the monomer pool of train 3 has onemember: propylene. Also, note that in this example, propylene is presentin all individual monomer pools (the pools of train 1, 2, and 3),therefore propylene is called a common member of all three monomerpools.

Unreduced reactor effluent refers to the whole effluent stream emergingfrom a reactor train, leg, or branch that has yet to undergo phaseseparation in one or more high pressure separators., i.e., it containsthe entire unreduced polymerization system as it emerges from thereactor train, leg, or branch. As opposed to unreduced reactoreffluents, reduced reactor effluents are the polymer-containing streamsderived from the unreduced reactor effluents. Reduced reactor effluentscontain less than the entire, unreduced polymerization system as itemerges from the reactor train, leg, or branch. In practical terms,reduced effluents are formed by separating and removing a part of themonomer and the optional inert solvent/diluent content in the form of amonomer-rich stream. In the practice of the present disclosure, theseparation of the said monomer-rich effluent and thus the reduction ofthe unreduced reactor effluent is performed by phase separation thatgenerates the said monomer-rich stream and the said reduced reactoreffluent, the latter of which in the form of a polymer-enriched stream.Stating it yet another way, the unreduced reactor effluent is the wholereactor effluent before any split of that effluent occurs, while areduced effluent is formed after at least one split in which some of thelight components (monomers and the optional inert solvents/diluents) areremoved from the whole, unreduced reactor effluent.

DETAILED DESCRIPTION

The monomer recycle processes disclosed herein operate in conjunctionwith the in-line fluid phase polymer blending processes described indetail in U.S. Patent Application No. 60/876,193 filed on Dec. 20, 2006,herein incorporated by reference in its entirety.

One form of the improved in-line process for blending polymers disclosedin U.S. Patent Application Ser. No. 60/876,193 includes (a) providingtwo or more reactor trains configured in parallel and two or morehigh-pressure separators fluidly connected to the two or more reactortrains configured in parallel; (b) contacting in the two or more reactortrains configured in parallel olefin monomers having two or more carbonatoms with: 1) one or more catalyst systems, 2) optional one or morecomonomers, 3) optional one or more scavengers, and 4) optional one ormore diluents or solvents, wherein at least one of the reactor trainsconfigured in parallel is at a temperature above the solid-fluidphase-transition temperature of the polymerization system and a pressureno lower than 10 MPa below the cloud point pressure of thepolymerization system and less than 1500 MPa, wherein at least one ofthe reactor trains includes an olefin monomer that has three or morecarbon atoms, wherein the polymerization system for each reactor trainis in its dense fluid state and comprises the olefin monomers, anycomonomer present, any diluent or solvent present, any scavengerpresent, and the polymer product, wherein the catalyst system for eachreactor train comprises one or more catalyst precursors, one or moreactivators, and optionally, one or more catalyst supports; (c) formingan unreduced reactor effluent including a homogenous fluid phasepolymer-monomer mixture in each parallel reactor train; (d) passing theunreduced reactor effluents from one or more but not from all of theparallel reactor trains through one or more high-pressure separators,maintaining the temperature and pressure within the one or morehigh-pressure separators above the solid-fluid phase transition pointbut below the cloud point pressure and temperature to form one or morefluid-fluid two-phase systems with each two-phase system comprising apolymer-rich phase and a monomer-enriched phase, and separating themonomer-rich phase from the polymer-rich phase in each of the one ormore high-pressure separators to form one or more separated monomer-richphases and one or more polymer-rich phases; (e) combining the one ormore polymer-enriched phases from the one or more high-pressureseparators of (d) with the one or more unreduced reactor effluents fromone or more parallel reactor trains to form a mixture of the one or morepolymer-enriched phases and the one or more unreduced reactor effluentsfrom the one or more parallel reactor trains to form a combined effluentstream that comprises the polymeric blend components from all parallelreactor trains; (f) passing the combined effluent stream of (e) intoanother high-pressure separator for product blending and product-feedseparation; (g) maintaining the temperature and pressure within theanother high pressure separator of (f) above the solid-fluid phasetransition point but below the cloud point pressure and temperature toform a fluid-fluid two-phase system comprising a polymer-rich blendphase and a monomer-rich phase; and (h) separating the monomer-richphase from the polymer-rich blend phase to form a polymer blend and aseparated monomer-rich phase. The polymer-rich blend phase is conveyedto a downstream finishing stage for further monomer stripping, dryingand/or pelletizing to form a polymer product blend. The fluid phasein-line blending process improves the resultant polymer blend byblending polymers in the presence of monomers as a solvent, which aidsblending by reducing viscosity in the blending step.

Another form of the improved in-line process for blending polymersdisclosed in U.S. Patent Application Ser. No. 60/876,193 includes: (a)providing two or more reactor trains configured in parallel and ahigh-pressure separator downstream fluidly connected to the two or morereactor trains configured in parallel; (b) contacting in the two or morereactor trains configured in parallel olefin monomers having two or morecarbon atoms with: 1) one or more catalyst systems, 2) optional one ormore comonomers, 3) optional one or more scavengers, and 4) optional oneor more diluents or solvents, wherein at least one of the reactor trainsconfigured in parallel is at a temperature above the solid-fluid phasetransition temperature of the polymerization system and a pressure nolower than 10 MPa below the cloud point pressure of the polymerizationsystem and less than 1500 MPa, wherein at least one of the reactortrains includes an olefin monomer that has three or more carbon atoms,wherein the polymerization system for each reactor train is in its densefluid state and comprises the olefin monomers, any comonomer present,any diluent or solvent present, any scavenger present, and the polymerproduct, wherein the catalyst system for each reactor train comprisesone or more catalyst precursors, one or more activators, and optionally,one or more catalyst supports; (c) forming a reactor effluent includinga homogeneous fluid phase polymer-monomer mixture in each parallelreactor train; (d) combining the reactor effluent comprising thehomogeneous fluid phase polymer-monomer mixture from each parallelreactor train to form a combined reactor effluent; (e) passing thecombined reactor effluent through the high-pressure separator forproduct blending and product-feed separation; (f) maintaining thetemperature and pressure within the high-pressure separator above thesolid-fluid phase transition point but below the cloud point pressureand temperature to form a fluid-fluid two-phase system comprising apolymer-rich blend phase and a monomer-rich phase; and (g) separatingthe monomer-rich phase from the polymer-rich blend phase to form apolymer blend and a separated monomer-rich phase. The polymer-rich blendphase is conveyed to a downstream finishing stage for further monomerstripping, drying and/or pelletizing to form a polymer product blend.The fluid phase in-line blending process improves the resultant polymerblend by blending polymers in the presence of monomers as a solvent,which aids blending by reducing viscosity in the blending step.

The novel recycle methods disclosed herein provide for simplifiedrecycle methods for the monomers emerge unconverted from the parallelreactor trains of the in-line fluid phase polymer blending processesdescribed in U.S. Patent Application No. 60/876,193 (will be referred toherein as the in-line fluid phase polymer blending process or thein-line polymer blending process). Particularly, the novel monomerrecycle methods are applicable for use with said fluid-phase in-linepolymer blending processes in which each monomer component fed to afirst group of one or more reactor trains of the said in-line blendingprocesses (G1) is also present in the feed of a second group of one ormore trains of the said in-line blending processes (G2) so that when themonomer pool of the said first group of trains (G1) is combined with themonomer pool of the second group of trains (G2), the said combinedmonomer pool and the monomer pool of the second group of trains (G2) arethe same. Stating it differently, when the effluents (or reducedeffluent streams derived from the effluents) of the said first group ofreactor trains (G1) are combined with the effluents of the said secondgroup of reactor trains (G2), the combined effluent stream essentiallycontains only monomers that are present in the feed of the said secondgroup of reactor trains (G2). Stating it yet another way, the effluents(or reduced effluent streams derived from the effluents) of the saidfirst group of reactor trains (G1) essentially do not introduce newmonomer components into the recycled effluents of said second group ofreactor trains (G2) when the effluent streams of G1 and G2 are combined.In a mathematical form, these conditions can be described as follows:

N(G1+G2)=N(G2) and N(G1)≦N(G2)

Where N(G1+G2) is the number of monomers in the combined monomer pool ofthe first and second group of reactor trains of the in-line fluid phasepolymer blending process; N(G1) and N(G2) are the number of monomers inthe monomer pool of the first (G1) and second (G2) group of reactortrains of the in-line fluid phase polymer blending process,respectively. The monomer pools present in the individual reactor trainsof G1 can be the same or different. However, the monomer pools presentin the individual reactor trains of G2 are always the same, although themonomer concentrations or monomer ratios may be different (but may alsobe the same). The number of reactor trains both in the first and in thesecond groups of reactor trains (G1 and G2) can be one or more than one.In practice, the number of reactor trains belonging to the first groupof reactor trains of the in-line fluid phase polymer blending processe(G1) can be one, two, three, four, five, or more. Similarly, the numberof reactor trains belonging to the second group of reactor trains of thein-line fluid phase polymer blending processes (G2) can also be one,two, three, four, five, or more. It should be understood that as allreactor trains of the in-line fluid phase polymer blending processesdisclosed herein, the one or more reactor trains of G1 are configured inparallel relative to the one or more reactor trains of G2. The G1 and G2reactor trains are also fluidly connected to one another. The totalnumber of reactors in the parallel reactor bank may be any number,although for economic reasons the number of reactors should bemaintained as low as the desired product grade slate and plant capacityallows. The optimum number of parallel reactor trains (also referred toas legs of the reactor bank) may be determined by standard chemicalengineering optimization methods well known in the art. Most typically,the polymerization-blending plant of the processes described in U.S.Patent Application No. 60/876,193 will have up to three polymerizationreactor trains or legs both in the first (G1) and in the second (G2)group of reactor trains. However, more than three reactor trains/legsper reactor group may be employed if the production of the targetproduct blends so requires.

When the above-stated conditions for the monomer pools are met in thein-line fluid phase polymer blending processes described in U.S. PatentApplication No. 60/876,193 filed on Dec. 20, 2006, the disclosed novelmonomer recycle methods provide for a simplified and thus advantagedmonomer recycle method for the monomer recycle streams. In allembodiments of the novel recycle processes disclosed herein, the monomerrecycle streams recovered from the product streams of G1 before mixingthem with any of the effluents of G2 are recycled to G1 while themonomer recycle streams recovered from the mixed polymer-containingstreams of G1 and G2 are recycled to G2. Since the mixed streams thatcontain monomers originated both from G1 and G2 are recycled to G2, thedisclosed novel monomer recycle methods also ensure that the monomercomponent recycle rates in the recycle stream originated from thecombined G1 and G2 product-containing streams and sent to G2 are nothigher than the desired monomer component flow rates in the compositefeed of G2. There are several methods available for achieving suchbalanced monomer recycles, which are disclosed below in the particularembodiments of the novel recycle processes. Note that meeting thismass-balance requirement in combination with the earlier-definedrequirement for the monomer pools of G1 and G2 allows for the recycle ofthe said monomer-rich streams to G2 unseparated and without building upexcess monomer inventories in G2.

In some embodiments of the novel recycle processes operating inconjunction with the in-line fluid phase polymer blending processesdescribed in detail in U.S. Patent Application No. 60/876,193, thecombined effluents of the first group of reactor trains (G1) delivermonomers at rates that are not higher than the combined purge rates ofthe corresponding monomers for the entire process plus the combinedconversion rates of the same monomers in the second group of reactortrains (G2). In these embodiments, the invention processes mix theunreduced effluents (i.e., without recovering monomers for recycle) ofthe first group of reactor trains (G1) with the effluents of the secondgroup of reactor trains (G2). The monomer-rich recycle streams emergingfrom and downstream of the mixing point of the polymer-containingeffluent streams of G1 and G2 are recycled to the second group ofreactor trains (G2). Optionally, some parts of the monomer recyclestreams are purged from the system to control the buildup of light inertcomponents, such as light alkanes (methane, ethane, propane, etc.),nitrogen, etc. Since these purge streams typically also containmonomers, the rate at which monomers are recycled to G2 are typicallysomewhat lower than the combined reactor effluent flow rates of thosemonomers. Note that since at least as much of the monomers common to thefirst (G1) and second (G2) groups of reactor trains is consumed in G2 orremoved in the optional purge streams as much brought in by theeffluents from G1, there is no excess monomer inventory buildup in G2even without separating and recovering some or all of the monomers fromthe reactor effluents or from the streams derived from the reactoreffluents before recycling them to G2. These embodiments are the simplerones among the novel recycle processes disclosed herein in that theyapply a single high-pressure separator to separate and recover amonomer-rich and a polymer-rich phase while blending the monomer andpolymer components originated from G1 and G2. Since this high-pressureseparator serves both as a separator and as a blender, it is alsoreferred to as separator-blender. The monomer-rich stream emerging fromthe separator-blender is recycled to G2 while the polymer-rich streamemerging from the separator-blender is sent to the finishing section forfurther monomer and lights removal. The monomer recycle streams emergingfrom and downstream of the separator-blender may be purged to controlundesirable heavy and light components, such as solvents, inert alkanes,N₂, excess catalyst killers, etc., before they are sent back to feed thesecond group of reactor trains. Note that in such embodiments, themonomers are recycled without recovering them individually before orafter the product streams of G1 and G2 are mixed saving substantialcapital and operating costs as compared to methods that separate andrecover the individual monomers from the recycle streams. Theseembodiments, however, require the combined unreduced effluents of thefirst group of reactor trains (G1) deliver monomers at rates not higherthan the combined purge rates of the corresponding monomers for theentire process plus the combined conversion rates of the correspondingmonomers in the second group of reactor trains (G2).

In other embodiments, the combined effluents of the earlier-describedfirst group of reactor trains (G1) of the in-line fluid phase polymerblending process carry one or more monomers at rates higher than thecombined purge rates of the corresponding monomers for the entireprocess plus the combined conversion rates of the same monomers in thesecond group of reactor trains (G2) of the in-line fluid phase polymerblending process. Stating it differently, one or more monomers arepresent in excess in the effluents of G1. As in all embodiments of thenovel recycle processes disclosed herein, the monomer recycle streamsrecovered after mixing all the product streams originated from G1 and G2are recycled to only one selected group of reactor trains, G2. Sincesuch recycle method would lead to a buildup of the excess monomers in G2in these embodiments, the excess monomer flows need to be removed byphase separation before mixing the product streams originated from G1and G2 and thus before they are recycled to G2. Therefore, inembodiments where the combined effluents of G1 carry one or moremonomers in excess, the disclosed processes remove the excess monomerstreams from the effluents of G1 prior to mixing the effluents with theproduct streams of G2. The removed excess monomer streams are recycledto their corresponding G1 reactor trains without separating andrecovering the individual monomers present in their essentially pureforms. The said excess monomer recovery is achieved by phase separationperformed in high-pressure separators as described in U.S. PatentApplication No. 60/876,193. It comprises the removal of enough monomersfrom the one or more effluents of the first group of reactor trains (G1)to ensure that after the reduction, the combined polymer-containingstreams originated from G1 bring in no excess monomer flows into therecycle streams sent to G2. It means that after removing the monomerexcesses, the combined polymer-containing product streams originatedfrom G1 deliver monomers to the blending point with the effluent streamsof G2 at rates that are not higher than the combined purge rates of thecorresponding monomers for the entire process plus the combinedconversion rates of the same monomers in the second group of reactortrains (G2). The monomer-rich recycle streams emerging from ordownstream of the mixing point of G1 and G2 product streams are recycledto G2. Consequently, at least a part of the monomer content of theproduct streams from G1 will be delivered to G2. Note that since noexcess monomer is brought in by streams mixed with the effluents of G2,no excess inventory of those monomers builds up in G2 even withoutrecovering some or all of the said monomers from the monomer-richrecycle streams recovered at or downstream of the said mixing pointbetween the product streams of G1 and G2. The separation and recovery ofexcess monomer flows from G1 for recycle to G1 can be performed by usingindividual reactor train effluents or by using combined effluent blendsof two or more reactor trains with same monomer pools. When G1 reactoreffluents are combined before excess monomer recovery for recycle to G1,the recycled monomer stream can be directed toward one or more of thereactor trains it originated from. As mentioned before, the separationprocesses disclosed herein achieve the removal of excess monomers fromthe reactor effluents of G1 by performing phase separations during whicha monomer-rich phase and a reduced polymer-containing (polymer-enriched)phase are recovered from the effluents of one or more reactor trainsbelonging to G1. Since these is phase separations are performed onstreams originated exclusively from reactor trains of G1, they will bereferred to as G1 phase separators, or G1 separators throughout thecurrent disclosure. Recycling the thus-separated one or moremonomer-rich streams to G1 ensures that no excess monomer stream isdirected toward G2. Naturally, since no foreign monomers are present inthese monomer-rich G1 recycle streams, there is no need for monomerseparation before their recycle to the appropriate reactor trains of G1.The reduced polymer-enriched streams along with other (unreduced)product streams of G1 in turn are mixed with the (unreduced) productstreams of G2 to ultimately yield an in-line polymer blend product. Thereduction of the effluents originated from the first group of reactortrains (G1), i.e. the recovery of excess monomer streams recycled totheir corresponding reactor trains belonging to G1, is controlled byadjusting the temperature and pressure of the phase separator so thatthe combined flow rate of each monomer component derived from G1 andblended with the product streams of G2 is not higher than the combinedpurge rate of each said monomer component for the entire process plusthe combined conversion rate of each said monomer component in G2. Theadjustment of the conditions of the phase separators for recovering theG1 monomer recycle streams can be readily performed by using standardchemical engineering techniques, including phase behavior measurements,well known in the art. The product-containing streams of G1 and G2 arecombined upstream of or in the high-pressure separator downstream of thelast G1 phase separator as described in described in U.S. PatentApplication No. 60/876,193. Since this high-pressure separator has adual function separating a monomer-rich stream from a polymer-richstream while blending the monomer and polymer components originated fromG1 and G2, it is also referred to as separator-blender. The monomer-richstream emerging from the separator-blender is recycled to G2. A part ofthis and other monomer recycle streams of the process may be purged tocontrol the buildup of undesirable heavy and light components, such asexcess solvents, catalyst killers, light alkanes, etc. The polymer-richstream emerging from the separator-blender is sent to the finishingsection of the in-line blending process disclosed in U.S. PatentApplication No. 60/876,193. The monomer-rich streams emerging downstreamof the separator-blender may be combined with the monomer-rich streamemerging from the separator-blender for recycle to G2. They may also bepartially or fully be purged from the system to control the buildup ofundesirable components. Note that monomer recoveries in the disclosedprocesses are achieved by phase separations as opposed to theconventional cryogenic distillation required for the separation ofmonomer recycle streams emerging downstream of the point where theproduct streams of G1 and G2 are mixed. Phase separations are simpler,require lower investment, and operate with less energy, making theprocesses disclosed herein advantageous over the conventional recyclemethods requiring costly distillative separation and recovery of theindividual monomers.

In some embodiments of the novel recycle processes operating inconjunction with the in-line fluid phase polymer blending processesdescribed in detail in U.S. Patent Application No. 60/876,193, thenumber of monomer components in any of monomer pools of the feeds to thefirst group of reactor trains (G1) may be one or more and the monomerpools of the reactor trains of G1 may be the same or different as longas the earlier-set N(G1+G2)=N(G2) and N(G1)≦N(G2) conditions for themonomer pools of the earlier-defined first (G1) and second (G2) groupsof reactor trains are met. For example, the reactor trains in G1 may allmake the same homopolymer (e.g., polypropylene, or polyethylene, orpolyhexene-1, etc.), or the same copolymer (e.g. ethylene-propylene orethylene-butene-1, etc. copolymers), or may make more than onehomopolymers (e.g., polyethylene and polypropylene) or more than onecopolymer, (e.g., ethylene-propylene and ethylene-butene-1, etc.,copolymers), or some may make a homopolymer and others may makecopolymers in any combination, as long as the monomer pools of the feedsto the reactor trains in G2 comprise all members of the combined monomerpools of the feeds to the reactor trains in G1. As in all embodiments ofthe novel recycle processes disclosed herein, the monomer recyclestreams recovered after mixing the product streams originated from G1and G2 are recycled to G2. The polymers made in G1 may have the same ordifferent average molecular weights, molecular weight distributions,melting and crystallization behaviors, melt viscosities, compositions,composition distributions, or combinations thereof. The second group ofreactor trains (G2) makes copolymers containing two or more monomersthat all have identical monomer pools and the feed monomer pool of saidsecond group of reactor trains (G2) comprises each monomer present inthe combined feed monomer pool of the first group of reactor trains(G1). The polymers made in the second group of reactor trains (G2) mayhave the same or different average molecular weights, molecular weightdistributions, melting and crystallization behaviors, average monomerconcentrations, composition distributions, or combination thereof. Theflow rate for each monomer is not more in the combined effluents of thefirst group of reactor trains (G1) than the combined purge rates of thecorresponding monomers for the entire process plus the combinedconversion rates of the corresponding monomers in the second group ofreactor trains (G2). The monomer-rich recycle streams emergingdownstream of the mixing point of the polymer-containing effluentstreams of G1 and G2 are recycled to the second group of reactor trains(G2). Optionally, some parts of the monomer recycle streams are purgedfrom the system to control the buildup of light inert components, suchas light alkanes (methane, ethane, propane, etc.), nitrogen, etc. Sincethese purge streams typically also contain monomers, the rate at whichmonomers are recycled to G2 are typically somewhat lower than thecombined reactor effluent flow rates of those monomers. Note that sinceat least as much of the monomers common to the first (G1) and second(G2) groups of reactor trains is consumed in G2 or removed in theoptional purge streams as much brought in by the effluents from G1,there is no excess monomer inventory buildup in G2 even withoutseparating and recovering some or all of the monomers from the reactoreffluents or from the streams derived from the reactor effluents beforerecycling them to G2. These embodiments are the simpler ones among thenovel recycle processes disclosed herein in that they apply a singlehigh-pressure separator to separate and recover a monomer-rich and apolymer-rich phase while blending the monomer and polymer componentsoriginated from G1 and G2. Since this high-pressure separator servesboth as a separator and a blender, it is also referred to asseparator-blender. The monomer-rich stream emerging from theseparator-blender is recycled to G2 while the polymer-rich streamemerging from the separator-blender is sent to the finishing section forfurther monomer and lights removal. The monomer recycle streams emergingfrom and downstream of the separator-blender may be purged to controlundesirable heavy and light components, such as solvents, inert alkanes,N₂, excess catalyst killers, etc., before they are sent back to feed thesecond group of reactor trains. Note that in such embodiments, themonomers are recycled without recovering them individually before orafter the product streams of G1 and G2 are mixed saving substantialcapital and operating costs as compared to methods that separate andrecover the individual monomers from the recycle streams. Theseembodiments, however, require the combined unreduced effluents of thefirst group of reactor trains (G1) deliver monomers at rates not higherthan the combined purge rates of the corresponding monomers for theentire process plus the combined conversion rates of the correspondingmonomers in the second group of reactor trains (G2).

In some embodiments of the novel recycle processes operating inconjunction with the in-line fluid phase polymer blending processesdescribed in detail in U.S. Patent Application No. 60/876,193, thenumber of monomer components in any of the monomer pools of the feeds tothe first group of reactor trains (G1) may be one or more and themonomer pools of the feeds to the trains of G1 may be the same ordifferent as long as the earlier-set N(G1+G2)=N(G2) and N(G1≦N(G2)conditions for the monomer pools of the earlier-defined first (G1) andsecond (G2) groups of reactor trains are met. For example, the reactortrains in G1 may all make the same homopolymer (e.g., polypropylene, orpolyethylene, or polyhexene-1, etc.), or the same copolymer (e.g.ethylene-propylene or ethylene-butene-1, etc. copolymers), or may makemore than one homopolymers (e.g., polyethylene and polypropylene) ormore than one copolymer, (e.g., ethylene-propylene andethylene-butene-1, etc., copolymers), or some may make a homopolymer andothers may make copolymers in any combination, as long as the monomerpools of the feeds to the reactor trains in G2 comprise all members ofthe combined monomer pools of the feeds to the reactor trains in G1. Asin all embodiments of the novel recycle processes disclosed herein, themonomer recycle streams recovered after mixing the product streamsoriginated from G1 and G2 are recycled to G2. The polymers made in G1may have the same or different average molecular weights, molecularweight distributions, melting and crystallization behaviors, meltviscosities, compositions, composition distributions, or combinationsthereof. The second group of reactor trains (G2) makes copolymerscontaining two or more monomers that all have identical monomer poolsand the feed monomer pool of said second group of reactor trains (G2)comprises each monomer present in the combined feed monomer pool of thefirst group of reactor trains (G1). The polymers made in the secondgroup of reactor trains (G2) may have the same or different averagemolecular weights, molecular weight distributions, melting andcrystallization behaviors, average monomer concentrations, compositiondistributions, or combination thereof. The combined effluents of theearlier-described first group of reactor trains (G1) carry one or moremonomers at rates higher than the combined purge rates of thecorresponding monomers for the entire process plus the combinedconversion rates of the same monomers in the second group of reactortrains (G2). Stating it differently, one or more monomers are present inexcess in the effluents of G1. As in all embodiments of the novelrecycle processes disclosed herein, the monomer recycle streamsrecovered after mixing all the product streams originated from G1 and G2are recycled to only one selected group of reactor trains, G2. Sincesuch recycle method would lead to a buildup of the excess monomers in G2in these embodiments, the excess monomer flows need to be removed byphase separation before mixing the product streams originated from G1and G2 and thus before they are recycled to G2. Therefore, inembodiments where the combined effluents of G1 carry one or moremonomers in excess, the disclosed processes remove the excess monomerstreams from the effluents of G1 prior to mixing the effluents with theproduct streams of G2. The removed excess monomer streams are recycledto their corresponding G1 reactor trains without separating andrecovering the individual monomers present in their essentially pureforms. The said excess monomer recovery is achieved by phase separationperformed in high-pressure separators as described in U.S. PatentApplication No. 60/876,193. It comprises the removal of enough monomersfrom the one or more effluents of the first group of reactor trains (G1)to ensure that after the reduction, the combined polymer-containingstreams originated from G1 bring in no excess monomer flows into therecycle streams sent to G2. It means that after removing the monomerexcesses, the combined polymer-containing product streams originatedfrom G1 deliver monomers to the blending point with the effluent streamsof G2 at rates that are not higher than the combined purge rates of thecorresponding monomers for the entire process plus the combinedconversion rates of the same monomers in the second group of reactortrains (G2). The monomer-rich recycle streams emerging downstream of themixing point of G1 and G2 product streams are recycled to G2.Consequently, at least a part of the monomer content of the productstreams from G1 will be delivered to G2. Note that since no excessmonomer is brought in by streams mixed with the effluents of G2, noexcess inventory of those monomers builds up in G2 even withoutrecovering some or all of the said monomers from the monomer-richrecycle streams recovered downstream of the said mixing point betweenthe product streams of G1 and G2. Typically, the high-pressure phaseseparators applied to reduce one or more effluent streams of reactortrains of G1 are used to treat the effluents of individual reactortrains of G1. However, when G1 has more than one trains with the samemonomer pools, their effluent streams may optionally be combined beforephase separation to remove the excess monomer flow. In no instances,however, may two or more reactor train effluent streams of G1 becombined to perform the removal of the excess monomer streams forrecycle to a reactor train in G1 if the monomer pools of the reactoreffluents involved were not identical. If the effluents of two or morereactor trains of G1 are combined before the removal of the excessmonomer flow, the recovered monomer stream can be recycled to any one ofthe originator reactor trains of G1 or can be split among them. Theproper choice can be determined by mass balance calculations well knownin the art of chemical engineering. As mentioned before, the separationprocesses disclosed herein achieve the removal of excess monomers fromthe reactor effluents of G1 by performing phase separations during whicha monomer-rich phase and a reduced polymer-containing (polymer-enriched)phase are recovered from the effluents of one or more reactor trainsbelonging to G1. Since these phase separations are performed on streamsoriginated exclusively from reactor trains of G1, they will be referredto as G1 phase separators, or G1 separators throughout the currentdisclosure. Recycling the thus-separated one or more monomer-richstreams to G1 ensures that no excess monomer stream is directed towardG2. Naturally, since no foreign monomers are present in thesemonomer-rich G1 recycle streams, there is no need for monomer separationbefore their recycle to the appropriate reactor trains of G1. Thereduced polymer-enriched streams along with other (unreduced) productstreams of G1 in turn are mixed with the (unreduced) product streams ofG2 to ultimately yield an in-line polymer blend product. The reductionof the effluents originated from the first group of reactor trains (G1),i.e. the recovery of excess monomer streams recycled to theircorresponding reactor trains belonging to G1, is controlled by adjustingthe temperature and pressure of the phase separator so that the combinedflow rate of each monomer component derived from G1 and blended with theproduct streams of G2 is not higher than the combined purge rate of eachsaid monomer component for the entire process plus the combinedconversion rate of each said monomer component in G2. The adjustment ofthe conditions of the phase separators for recovering the G1 monomerrecycle streams can be readily performed by using standard chemicalengineering techniques, including phase behavior measurements, wellknown in the art. The product-containing streams of G1 and G2 arecombined upstream of or in the high-pressure separator downstream of thelast G1 phase separator as described in described in U.S. PatentApplication No. 60/876,193. Since this high-pressure separator has adual function separating a monomer-rich stream from a polymer-richstream while blending the monomer and polymer components originated fromG1 and G2, it is also referred to as separator-blender. The monomer-richstream emerging from the separator-blender is recycled to G2. A part ofthis and other monomer recycle streams of the process may be purged tocontrol the buildup of undesirable heavy and light components, such asexcess solvents, catalyst killers, light alkanes, etc. The polymer-richstream emerging from the separator-blender is sent to the finishingsection of the in-line blending process disclosed in U.S. PatentApplication No. 60/876,193. The monomer-rich streams emerging downstreamof the separator-blender may be combined with the monomer-rich streamemerging from the separator-blender for recycle to G2. They may also bepartially or fully be purged from the system to control the buildup ofundesirable components. Note that this embodiment corrects for themonomer excesses present in the combined effluent streams of G1 and G2by partially recovering the excess monomers by phase separation prior toblending the polymer-containing product streams of G1 and G2 in theseparator-blender of the process described in U.S. Patent ApplicationNo. 60/876,193. Note also that the said separator-blender of the processdescribed in U.S. Patent Application No. 60/876,193 is the firsthigh-pressure separator among the serially connected separators thatcombine all the polymer-containing effluents of G1 and G2. It is alsoimportant to emphasize that the monomer-rich effluents originated fromand downstream of the said separator-blender are recycled to the secondgroup of trains (G2) and not to the first group of reactor trains (G1).The monomer-rich streams recovered by the single-line G1 separatorsupstream of the said separator-blender, on the other hand, are alwaysrecycled directly to their own reactor trains belonging to G1 (alsowithout monomer separation). The destination of the recycled monomerstream recovered in a high-pressure separator, therefore, divides thehigh-pressure separators into two groups: one sends back the recoveredmonomer-rich streams to G1, while the other one to G2. The first amongthe serially connected separators of the novel monomer recycle processesdescribed herein that send their monomer-rich recycle streams to G2 isalso the first that combines all polymer-containing product streams.Since it combines the functions of polymer-monomer separation andpolymer-polymer blending, it is also referred to as separator-blender.We need to point out that in our descriptions the separator-blenderdesignation is reserved to this one unique, readily identifiablehigh-pressure separator. Therefore, the earlier mentioned optionalseparators sending monomer-enriched streams to reactor trains in G1 arenever called separator-blenders in our descriptions even if theirfunction is to combine two or more reactor train effluents. Instead,they are referred to as G1 phase separators or G1 separators. It isimportant to point out that monomer recoveries in the disclosedprocesses are achieved by phase separations as opposed to theconventional cryogenic distillation required for the separation ofmonomer recycle streams emerging downstream of the point where theproduct streams of G1 and G2 are mixed. Phase separations are simpler,require lower investment, and operate with less energy, making theprocesses disclosed herein advantageous over the conventional recyclemethods requiring costly distillative separation and recovery of theindividual monomers.

In other embodiments of the novel recycle processes operating inconjunction with the in-line fluid phase polymer blending processesdescribed in detail in U.S. Patent Application No. 60/876,193, the feedmonomer pool of each reactor train of the first group of reactor trains(G1) has only one monomer member. The monomers fed to the differenttrains of the first group of reactor trains (G1) may be the same ordifferent. In these embodiments, the members of the first group ofreactor trains (G1) are also referred to as homopolymer reactor trainsor homopolymer trains or homopolymerization reactor trains orhomopolymerization trains. The homopolymers made in the homopolymertrains can have the same or different average molecular weights,molecular weight distributions, melting and crystallization behaviors,melt viscosities, compositions, or combination thereof. The second groupof reactor trains (G2), the members of which are also referred to ascopolymer reactor trains or copolymer trains or copolymerization reactortrains or copolymerization trains, makes copolymers that all compriseidentical monomer pools (i.e., composed of the exact same monomers) andthe monomer pool of said copolymerization trains (G2) comprises eachmonomer present in the combined monomer pool of the homopolymer trains(G1). The polymers made in the copolymer trains are composed of the samemonomer units and may have the same or different average molecularweights, molecular weight distributions, melting and crystallizationbehaviors, melt viscosities, average monomer concentrations, compositiondistributions, or combination thereof. For example, the homopolymercould be polypropylene and the copolymer could be ethylene-propylene orpropylene-butene-1, or propylene-hexene-1, or propylene-octene-1, orpropylene-decene-1 copolymer, etc. The flow rate for each monomer is notmore in the combined effluents of the first group of reactor trains (G1)than the combined purge rates of the corresponding monomers for theentire process plus the combined conversion rates of the correspondingmonomers in the second group of reactor trains (G2). The monomer-richrecycle streams emerging downstream of the mixing point of thepolymer-containing effluent streams of G1 and G2 are recycled to thesecond group of reactor trains (G2). Optionally, some parts of themonomer recycle streams are purged from the system to control thebuildup of light inert components, such as light alkanes (methane,ethane, propane, etc.), nitrogen, etc. Since these purge streamstypically also contain monomers, the rate at which monomers are recycledto G2 are typically somewhat lower than the combined reactor effluentflow rates of those monomers. Note that since at least as much of themonomers common to the first (G1) and second (G2) groups of reactortrains is consumed in G2 or removed in the optional purge streams asmuch brought in by the effluents from G1, there is no excess monomerinventory buildup in G2 even without separating and recovering some orall of the monomers from the reactor effluents or from the streamsderived from the reactor effluents before recycling them to G2. Theseembodiments are the simpler ones among the novel recycle processesdisclosed herein in that they apply a single high-pressure separator toseparate and recover a monomer-rich and a polymer-rich phase whileblending the monomer and polymer components originated from G1 and G2.Since this high-pressure separator serves both as a separator and ablender, it is also referred to as separator-blender. The monomer-richstream emerging from the separator-blender is recycled to G2 while thepolymer-rich stream emerging from the separator-blender is sent to thefinishing section for further monomer and lights removal. The monomerrecycle streams emerging from and downstream of the separator-blendermay be purged to control undesirable heavy and light components, such assolvents, inert alkanes, N₂, excess catalyst killers, etc., before theyare sent back to feed the second group of reactor trains. Note that insuch embodiments, the monomers are recycled without recovering themindividually before or after the product streams of G1 and G2 are mixedsaving substantial capital and operating costs as compared to methodsthat separate and recover the individual monomers from the recyclestreams. These embodiments, however, require the combined unreducedeffluents of the first group of reactor trains (G1) deliver monomers atrates not higher than the combined purge rates of the correspondingmonomers for the entire process plus the combined conversion rates ofthe corresponding monomers in the second group of reactor trains (G2).

In other embodiments of the novel recycle processes operating inconjunction with the in-line fluid phase polymer blending processesdescribed in detail in U.S. Patent Application No. 60/876,193, the feedmonomer pool of each reactor train of the first group of reactor trains(G1) has only one monomer member. The monomers fed to the differenttrains of the first group of reactor trains (G1) may be the same ordifferent. In these embodiments, the members of the first group ofreactor trains (G1) are also referred to as homopolymer reactor trainsor homopolymer trains or homopolymerization reactor trains orhomopolymerization trains. The homopolymers made in the homopolymertrains can have the same or different average molecular weights,molecular weight distributions, melting and crystallization behaviors,melt viscosities, compositions, or combination thereof. The second groupof reactor trains (G2), the members of which are also referred to ascopolymer reactor trains or copolymer trains or copolymerization reactortrains or copolymerization trains, makes copolymers that all compriseidentical monomer pools (i.e., composed of the exact same monomers) andthe monomer pool of said copolymerization trains (G2) comprises eachmonomer present in the combined monomer pool of the homopolymer trains(G1). The polymers made in the copolymer trains are composed of the samemonomer units and may have the same or different average molecularweights, molecular weight distributions, melting and crystallizationbehaviors, melt viscosities, average monomer concentrations, compositiondistributions, or combination thereof. For example, the homopolymercould be polypropylene and the copolymer could be ethylene-propylene orpropylene-butene-1, or propylene-hexene-1, or propylene-octene-1, orpropylene-decene-1 copolymer, etc. The combined effluents of theearlier-described first group of reactor trains (G1) carry one or moremonomers at rates higher than the combined purge rates of thecorresponding monomers for the entire process plus the combinedconversion rates of the same monomers in the second group of reactortrains (G2). Stating it differently, one or more monomers are present inexcess in the effluents of G1. As in all embodiments of the novelrecycle processes disclosed herein, the monomer recycle streamsrecovered after mixing all the product streams originated from G1 and G2are recycled to only one selected group of reactor trains, G2. Sincesuch recycle method would lead to a buildup of the excess monomers in G2in these embodiments, the excess monomer flows need to be removed byphase separation before mixing the product streams originated from G1and G2 and thus before they are recycled to G2. Therefore, inembodiments where the combined effluents of G1 carry one or moremonomers in excess, the disclosed processes remove the excess monomerstreams from the effluents of G1 prior to mixing the effluents with theproduct streams of G2. The removed excess monomer streams are recycledto their corresponding G1 reactor trains without separating andrecovering the individual monomers present in their essentially pureforms. The said excess monomer recovery is achieved by phase separationperformed in high-pressure separators as described in U.S. PatentApplication No. 60/876,193. It comprises the removal of enough monomersfrom the one or more effluents of the first group of reactor trains (G1)to ensure that after the reduction, the combined polymer-containingstreams originated from G1 bring in no excess monomer flows into therecycle streams sent to G2. It means that after removing the monomerexcesses, the combined polymer-containing product streams originatedfrom G1 deliver monomers to the blending point with the effluent streamsof G2 at rates that are not higher than the combined purge rates of thecorresponding monomers for the entire process plus the combinedconversion rates of the same monomers in the second group of reactortrains (G2). The monomer-rich recycle streams emerging downstream of themixing point of G1 and G2 product streams are recycled to G2.Consequently, at least a part of the monomer content of the productstreams from G1 will be delivered to G2. Note that since no excessmonomer is brought in by streams mixed with the effluents of G2, noexcess inventory of those monomers builds up in G2 even withoutrecovering some or all of the said monomers from the monomer-richrecycle streams recovered downstream of the said mixing point betweenthe product streams of G1 and G2. Typically, the high-pressure phaseseparators applied to reduce one or more effluent streams of reactortrains of G1 are used to treat the effluents of individual reactortrains of G1. However, when G1 has more than one trains with the samemonomer pools, their effluent streams may optionally be combined beforephase separation to remove the excess monomer flow. In no instances,however, may two or more reactor train effluent streams of G1 becombined to perform the removal of the excess monomer streams forrecycle to a reactor train in G1 if the monomer pools of the reactoreffluents involved were not identical. If the effluents of two or morereactor trains of G1 are combined before the removal of the excessmonomer flow, the recovered monomer stream can be recycled to any one ofthe originator reactor trains of G1 or can be split among them. Theproper choice can be determined by mass balance calculations well knownin the art of chemical engineering. As mentioned before, the separationprocesses disclosed herein achieve the removal of excess monomers fromthe reactor effluents of G1 by performing phase separations during whicha monomer-rich phase and a reduced polymer-containing (polymer-enriched)phase are recovered from the effluents of one or more reactor trainsbelonging to G1. Since these phase separations are performed on streamsoriginated exclusively from reactor trains of G1, they will be referredto as G1 phase separators, or G1 separators throughout the currentdisclosure. Recycling the thus-separated one or more monomer-richstreams to G1 ensures that no excess monomer stream is directed towardG2. Naturally, since no foreign monomers are present in thesemonomer-rich G1 recycle streams, there is no need for monomer separationbefore their recycle to the appropriate reactor trains of G1. Thereduced polymer-enriched streams along with other (unreduced) productstreams of G1 in turn are mixed with the (unreduced) product streams ofG2 to ultimately yield an in-line polymer blend product. The reductionof the effluents originated from the first group of reactor trains (G1),i.e. the recovery of excess monomer streams recycled to theircorresponding reactor trains belonging to G1, is controlled by adjustingthe temperature and pressure of the phase separator so that the combinedflow rate of each monomer component derived from G1 and blended with theproduct streams of G2 is not higher than the combined purge rate of eachsaid monomer component for the entire process plus the combinedconversion rate of each said monomer component in G2. The adjustment ofthe conditions of the phase separators for recovering the G1 monomerrecycle streams can be readily performed by using standard chemicalengineering techniques, including phase behavior measurements, wellknown in the art. The product-containing streams of G1 and G2 arecombined upstream of or in the high-pressure separator downstream of thelast G1 phase separator as described in described in U.S. PatentApplication No. 60/876,193. Since this high-pressure separator has adual function separating a monomer-rich stream from a polymer-richstream while blending the monomer and polymer components originated fromG1 and G2, it is also referred to as separator-blender. The monomer-richstream emerging from the separator-blender is recycled to G2. A part ofthis and other monomer recycle streams of the process may be purged tocontrol the buildup of undesirable heavy and light components, such asexcess solvents, catalyst killers, light alkanes, etc. The polymer-richstream emerging from the separator-blender is sent to the finishingsection of the in-line blending process disclosed in U.S. PatentApplication No. 60/876,193. The monomer-rich streams emerging downstreamof the separator-blender may be combined with the monomer-rich streamemerging from the separator-blender for recycle to G2. They may also bepartially or fully be purged from the system to control the buildup ofundesirable components. Note that this embodiment corrects for themonomer excesses present in the combined effluent streams of G1 and G2by partially recovering the excess monomers by phase separation prior toblending the polymer-containing product streams of G1 and G2 in theseparator-blender of the process described in U.S. Patent ApplicationNo. 60/876,193. Note also that the said separator-blender of the processdescribed in U.S. Patent Application No. 60/876,193 is the firsthigh-pressure separator among the serially connected separators thatcombine all the polymer-containing effluents of G1 and G2. It is alsoimportant to emphasize that the monomer-rich effluents originated fromand downstream of the said separator-blender are recycled to the secondgroup of trains (G2) and not to the first group of reactor trains (G1).The monomer-rich streams recovered by the single-line G1 separatorsupstream of the said separator-blender, on the other hand, are alwaysrecycled directly to their own reactor trains belonging to G1 (alsowithout monomer separation). The destination of the recycled monomerstream recovered in a high-pressure separator, therefore, divides thehigh-pressure separators into two groups: one sends back the recoveredmonomer-rich streams to G1, while the other one to G2. The first amongthe serially connected separators of the novel monomer recycle processesdescribed herein that send their monomer-rich recycle streams to G2 isalso the first that combines all polymer-containing product streams.Since it combines the functions of polymer-monomer separation andpolymer-polymer blending, it is also referred to as separator-blender.We need to point out that in our descriptions the separator-blenderdesignation is reserved to this one unique, readily identifiablehigh-pressure separator. Therefore, the earlier mentioned optionalseparators sending monomer-enriched streams to reactor trains in G1 arenever called separator-blenders in our descriptions even if theirfunction is to combine two or more reactor train effluents. Instead,they are referred to as G1 phase separators or G1 separators. It isimportant to point out that monomer recoveries in the disclosedprocesses are achieved by phase separations as opposed to theconventional cryogenic distillation required for the separation ofmonomer recycle streams emerging downstream of the point where theproduct streams of G1 and G2 are mixed. Phase separations are simpler,require lower investment, and operate with less energy, making theprocesses disclosed herein advantageous over the conventional recyclemethods requiring costly distillative separation and recovery of theindividual monomers.

In some other embodiments of the novel recycle processes operating inconjunction with the in-line fluid phase polymer blending processesdescribed in detail in U.S. Patent Application No. 60/876,193, the feedmonomer pools of the first group of reactor trains (G1) have one memberand they are the same in each monomer pool. The first group of reactortrains, the members of which are also referred to as homopolymer reactortrains or homopolymer trains or homopolymerization reactor trains orhomopolymerization trains. The homopolymers made in the homopolymertrains can have the same or different average molecular weights,molecular weight distributions, or combination thereof. The second groupof reactor trains (G2), the members of which are also referred to ascopolymer reactor trains or copolymer trains or copolymerization reactortrains or copolymerization trains, makes copolymers that all compriseidentical monomer pools (i.e., composed of the exact same monomers) andthe monomer pool of said copolymerization trains (G2) comprises themonomer present in the homopolymer trains (G1). The polymers made in thecopolymerization trains are composed of the same monomers and may havethe same or different average molecular weight, molecular weightdistribution, average monomer concentration, composition distribution,or combination thereof. For example, the homopolymer could bepolypropylene and the copolymer could be ethylene-propylene orpropylene-butene-1, or propylene-hexene-1, or propylene-octene-1, orpropylene-decene-1 copolymer, etc. The flow rate for the monomer is notmore in the combined effluents of the first group of reactor trains (G1)than the combined purge rates of the corresponding monomer for theentire process plus the combined conversion rates of the correspondingmonomer in the second group of reactor trains (G2). As in eachembodiment of the novel recycle processes disclosed herein, themonomer-rich recycle streams emerging downstream of the mixing point ofthe polymer-containing effluent streams of G1 and G2 are recycled to thesecond group of reactor trains (G2). Optionally, some parts of themonomer recycle streams are purged from the system to control thebuildup of light inert components, such as light alkanes (methane,ethane, propane, etc.), nitrogen, etc. Since these purge streamstypically also contain monomers, the rate at which monomers are recycledto G2 are typically somewhat lower than the combined reactor effluentflow rates of those monomers. Note that since at least as much of themonomer common to the first (G1) and second (G2) groups of reactortrains is consumed in G2 or removed in the optional purge streams asmuch brought in by the effluents from G1, there is no excess monomerinventory buildup in G2 even without separating and recovering some orall of the monomer from the reactor effluents or from the streamsderived from the reactor effluents before recycling them to G2. Theseembodiments are the simpler ones among the novel recycle processesdisclosed herein in that they apply a single high-pressure separator toseparate and recover a monomer-rich and a polymer-rich phase whileblending the monomer and polymer components originated from G1 and G2.Since this high-pressure separator serves both as a separator and ablender, it is also referred to as separator-blender. The monomer-richstream emerging from the separator-blender is recycled to G2 while thepolymer-rich stream emerging from the separator-blender is sent to thefinishing section for further monomer and lights removal. The monomerrecycle streams emerging from and downstream of the separator-blendermay be purged to control undesirable heavy and light components, such assolvents, inert alkanes, N₂, excess catalyst killers, etc., before theyare sent back to feed the second group of reactor trains. Note that insuch embodiments, the monomers are recycled without recovering themindividually before or after the product streams of G1 and G2 are mixedsaving substantial capital and operating costs as compared to methodsthat separate and recover the individual monomers from the recyclestreams. These embodiments, however, require the combined unreducedeffluents of the first group of reactor trains (G1) deliver the monomerat rates not higher than the combined purge rates of the correspondingmonomer for the entire process plus the combined conversion rates of thecorresponding monomer in the second group of reactor trains (G2).

In some other embodiments of the novel recycle processes operating inconjunction with the in-line fluid phase polymer blending processesdescribed in detail in U.S. Patent Application No. 60/876,193, the feedmonomer pools of the first group of reactor trains (G1) have one memberand they are the same in each monomer pool. The first group of reactortrains, the members of which are also referred to as homopolymer reactortrains or homopolymer trains or homopolymerization reactor trains orhomopolymerization trains. The homopolymers made in the homopolymertrains can have the same or different average molecular weights,molecular weight distributions, or combination thereof. The second groupof reactor trains (G2), the members of which are also referred to ascopolymer reactor trains or copolymer trains or copolymerization reactortrains or copolymerization trains, makes copolymers that all compriseidentical monomer pools (i.e., composed of the exact same monomers) andthe monomer pool of said copolymerization trains (G2) comprises themonomer present in the homopolymer trains (G1). The polymers made in thecopolymerization trains are composed of the same monomers and may havethe same or different average molecular weight, molecular weightdistribution, average monomer concentration, composition distribution,or combination thereof. For example, the homopolymer could bepolypropylene and the copolymer could be ethylene-propylene orpropylene-butene-1, or propylene-hexene-1, or propylene-octene-1, orpropylene-decene-1 copolymer, etc. The combined effluents of theearlier-described first group of reactor trains (G1) carry the monomerat rates higher than the combined purge rates of the correspondingmonomer for the entire process plus the combined conversion rates of thesame monomer in the second group of reactor trains (G2). Stating itdifferently, the monomer is present in excess in the effluents of G1. Asin all embodiments of the novel recycle processes disclosed herein, themonomer recycle streams recovered after mixing all the product streamsoriginated from G1 and G2 are recycled to only one selected group ofreactor trains, G2. Since such recycle method would lead to a buildup ofthe excess monomers in G2 in these embodiments, the excess monomer flowneed to be removed by phase separation before mixing the product streamsoriginated from G1 and G2 and thus before they are recycled to G2.Therefore, in embodiments where the combined effluents of G1 carry themonomer in excess, the disclosed processes remove the excess monomerstream from the effluents of G1 prior to mixing the effluents with theproduct streams of G2. The removed excess monomer stream is recycled totheir corresponding G1 reactor trains without separating and recoveringthe individual monomer present in its essentially pure forms. The saidexcess monomer recovery is achieved by phase separation performed inhigh-pressure separators as described in U.S. Patent Application No.60/876,193. It comprises the removal of enough monomer from the one ormore effluents of the first group of reactor trains (G1) to ensure thatafter the reduction, the combined polymer-containing streams originatedfrom G1 bring in no excess monomer flows into the recycle streams sentto G2. It means that after removing the monomer excess, the combinedpolymer-containing product streams originated from G1 deliver monomer tothe blending point with the effluent streams of G2 at a rate not higherthan the combined purge rates of the corresponding monomer for theentire process plus the combined conversion rates of the same monomer inthe second group of reactor trains (G2). The monomer-rich recyclestreams emerging downstream of the mixing point of G1 and G2 productstreams are recycled to G2. Consequently, at least a part of the monomercontent of the product streams from G1 will be delivered to G2. Notethat since no excess monomer is brought in by streams mixed with theeffluents of G2, no excess inventory of those monomers builds up in G2even without recovering some or all of the said monomer from themonomer-rich recycle streams recovered downstream of the said mixingpoint between the product streams of G1 and G2. Typically, thehigh-pressure phase separators applied to reduce one or more effluentstreams of reactor trains of G1 are used to treat the effluents ofindividual reactor trains of G1. However, when G1 has more than onetrains, their effluent streams may optionally be combined before phaseseparation to remove the excess monomer flow. If the effluents of two ormore reactor trains of G1 are combined before the removal of the excessmonomer flow, the recovered monomer stream can be recycled to any one ofthe reactor trains of G1 or can be split among them. The proper choicecan be determined by mass balance calculations well known in the art ofchemical engineering. As mentioned before, the separation processesdisclosed herein achieve the removal of excess monomers from the reactoreffluents of G1 by performing phase separations during which amonomer-rich phase and a reduced polymer-containing (polymer-enriched)phase are recovered from the effluents of one or more reactor trainsbelonging to G1. Since these phase separations are performed on streamsoriginated exclusively from reactor trains of G1, they will be referredto as G1 phase separators, or G1 separators throughout the currentdisclosure. Recycling the thus-separated one or more monomer-richstreams to G1 ensures that no excess monomer stream is directed towardG2. Naturally, since no foreign monomer is present in these monomer-richG1 recycle streams, there is no need for monomer separation before theirrecycle to the appropriate reactor trains of G1. The reducedpolymer-enriched streams along with other (unreduced) product streams ofG1 in turn are mixed with the (unreduced) product streams of G2 toultimately yield an in-line polymer blend product. The reduction of theeffluents originated from the first group of reactor trains (G1), i.e.the recovery of excess monomer stream recycled to their correspondingreactor trains belonging to G1, is controlled by adjusting thetemperature and pressure of the phase separator so that the combinedflow rate of the monomer component derived from G1 and blended with theproduct streams of G2 is not higher than the combined purge rate of thesaid monomer component for the entire process plus the combinedconversion rate of the said monomer component in G2. The adjustment ofthe conditions of the phase separators for recovering the G1 monomerrecycle streams can be readily performed by using standard chemicalengineering techniques, including phase behavior measurements, wellknown in the art. The product-containing streams of G1 and G2 arecombined upstream of or in the high-pressure separator downstream of thelast G1 phase separator as described in described in U.S. PatentApplication No. 60/876,193. Since this high-pressure separator has adual function separating a monomer-rich stream from a polymer-richstream while blending the monomer and polymer components originated fromG1 and G2, it is also referred to as separator-blender. The monomer-richstream emerging from the separator-blender is recycled to G2. A part ofthis and other monomer recycle streams of the process may be purged tocontrol the buildup of undesirable heavy and light components, such asexcess solvents, catalyst killers, light alkanes, etc. The polymer-richstream emerging from the separator-blender is sent to the finishingsection of the in-line blending process disclosed in U.S. PatentApplication No. 60/876,193. The monomer-rich streams emerging downstreamof the separator-blender may be combined with the monomer-rich streamemerging from the separator-blender for recycle to G2. They may also bepartially or fully be purged from the system to control the buildup ofundesirable components. Note that this embodiment corrects for themonomer excesses present in the combined effluent streams of G1 and G2by partially recovering the excess monomers by phase separation prior toblending the polymer-containing product streams of G1 and G2 in theseparator-blender of the process described in U.S. Patent ApplicationNo. 60/876,193. Note also that the said separator-blender of the processdescribed in U.S. Patent Application No. 60/876,193 is the firsthigh-pressure separator among the serially connected separators thatcombine all the polymer-containing effluents of G1 and G2. It is alsoimportant to emphasize that the monomer-rich effluents originated fromand downstream of the said separator-blender are recycled to the secondgroup of trains (G2) and not to the first group of reactor trains (G1).The monomer-rich streams recovered by the single-line G1 separatorsupstream of the said separator-blender, on the other hand, are alwaysrecycled directly to their own reactor trains belonging to G1 (alsowithout monomer separation). The destination of the recycled monomerstream recovered in a high-pressure separator, therefore, divides thehigh-pressure separators into two groups: one sends back the recoveredmonomer-rich streams to G1, while the other one to G2. The first amongthe serially connected separators of the novel monomer recycle processesdescribed herein that send their monomer-rich recycle streams to G2 isalso the first that combines all polymer-containing product streams.Since it combines the functions of polymer-monomer separation andpolymer-polymer blending, it is also referred to as separator-blender.We need to point out that in our descriptions the separator-blenderdesignation is reserved to this one unique, readily identifiablehigh-pressure separator. Therefore, the earlier mentioned optionalseparators sending monomer-enriched streams to reactor trains in G1 arenever called separator-blenders in our descriptions even if theirfunction is to combine two or more reactor train effluents. Instead,they are referred to as G1 phase separators or G1 separators. It isimportant to point out that monomer recoveries in the disclosedprocesses are achieved by phase separations as opposed to theconventional cryogenic distillation required for the separation ofmonomer recycle streams emerging downstream of the point where theproduct streams of G1 and G2 are mixed. Phase separations are simpler,require lower investment, and operate with less energy, making theprocesses disclosed herein advantageous over the conventional recyclemethods requiring costly distillative separation and recovery of theindividual monomers.

In another embodiment of the novel recycle processes operating inconjunction with the in-line fluid phase polymer blending processesdescribed in detail in U.S. Patent Application No. 60/876,193, there isone homopolymer train (G1) and one copolymer train (G2), i.e., both G1and G2 have one reactor train each and the feed monomer pool of G1 hasone monomer. The monomer pool of the copolymer comprises the monomer ofthe homopolymer. As stated earlier, the number of different monomers forthe combined process yielding the polymer blend is equal to the numberof monomer components used in making the copolymer component of thein-line product blend. For example, the homopolymer could bepolypropylene and the copolymer could be ethylene-propylene orpropylene-butene-1, or propylene-hexene-1, or propylene-octene-1, orpropylene-decene-1 copolymer, etc. The present disclosure does notcomprehend polymerization processes that would make blends ofhomopolymers and copolymers where there is not a single monomer commonto the homopolymer and copolymer (i.e. blends of polypropylene andethylene-butene-1 copolymer or polyethylene and propylene-hexene-1copolymer). Hence the monomer recycle process for use with fluid phasein-line polymer blending disclosed herein, the number of differentmonomers in the overall process is equal to the number of monomers inthe copolymerization reactor train. The flow rate for the monomer is notmore in the first reactor train (G1) than the combined purge rates ofthe corresponding monomer for the entire process plus the combinedconversion rates of the corresponding monomer in the second reactortrain (G2). As in each embodiment of the novel recycle processesdisclosed herein, the monomer-rich recycle streams emerging downstreamof the mixing point of the polymer-containing effluent streams of G1 andG2 are recycled to the second group of reactor trains (G2). Optionally,some parts of the monomer recycle streams are purged from the system tocontrol the buildup of light inert components, such as light alkanes(methane, ethane, propane, etc.), nitrogen, etc. Since these purgestreams typically also contain monomers, the rate at which monomers arerecycled to G2 are typically somewhat lower than the combined reactoreffluent flow rates of those monomers. Note that since at least as muchof the monomer common to the first (G1) and second (G2) groups ofreactor trains is consumed in G2 or removed in the optional purgestreams as much brought in by the effluents from G1, there is no excessmonomer inventory buildup in G2 even without separating and recoveringsome or all of the monomer from the reactor effluents or from thestreams derived from the reactor effluents before recycling them to G2.These embodiments are the simpler ones among the novel recycle processesdisclosed herein in that they apply a single high-pressure separator toseparate and recover a monomer-rich and a polymer-rich phase whileblending the monomer and polymer components originated from G1 and G2.Since this high-pressure separator serves both as a separator and ablender, it is also referred to as separator-blender. The monomer-richstream emerging from the separator-blender is recycled to G2 while thepolymer-rich stream emerging from the separator-blender is sent to thefinishing section for further monomer and lights removal. The monomerrecycle streams emerging from and downstream of the separator-blendermay be purged to control undesirable heavy and light components, such assolvents, inert alkanes, N₂, excess catalyst killers, etc., before theyare sent back to feed the second group of reactor trains. Note that insuch embodiments, the monomers are recycled without recovering themindividually before or after the product streams of G1 and G2 are mixedsaving substantial capital and operating costs as compared to methodsthat separate and recover the individual monomers from the recyclestreams. These embodiments, however, require the unreduced effluent ofthe homopolymerization train (G1) delivers the monomer at a rate nothigher than the combined purge rates of the corresponding monomer forthe entire process plus the combined conversion rate of thecorresponding monomer in the copolymerization reactor train (G2).

In another embodiment of the novel recycle processes operating inconjunction with the in-line fluid phase polymer blending processesdescribed in detail in U.S. Patent Application No. 60/876,193, there isone homopolymer train (G1) and one copolymer train (G2), i.e., both G1and G2 have one reactor train each and the feed monomer pool of G1 hasone monomer. The monomer pool of the copolymer comprises the monomer ofthe homopolymer. As stated earlier, the number of different monomers forthe combined process yielding the polymer blend is equal to the numberof monomer components used in making the copolymer component of thein-line product blend. For example, the homopolymer could bepolypropylene and the copolymer could be ethylene-propylene orpropylene-butene-1, or propylene-hexene-1, or propylene-octene-1, orpropylene-decene-1 copolymer, etc. The present disclosure does notcomprehend polymerization processes that would make blends ofhomopolymers and copolymers where there is not a single monomer commonto the homopolymer and copolymer (i.e. blends of polypropylene andethylene-butene-1 copolymer or polyethylene and propylene-hexene-1copolymer). Hence the monomer recycle process for use with fluid phasein-line polymer blending disclosed herein, the number of differentmonomers in the overall process is equal to the number of monomers inthe copolymerization reactor train. The effluent of thehomopolymerization reactor train (G1) carries the monomer at rateshigher than the combined purge rates of the corresponding monomer forthe entire process plus the combined conversion rate of the same monomerin the copolymerization train (G2). Stating it differently, the monomeris present in excess in the effluent of G1. As in all embodiments of thenovel recycle processes disclosed herein, the monomer recycle streamsrecovered after mixing the product streams originated from G1 and G2 arerecycled to only one selected group of reactor trains, G2. Since suchrecycle method would lead to a buildup of the excess monomers in G2 inthese embodiments, the excess monomer flow need to be removed by phaseseparation before mixing the product streams originated from G1 and G2and thus before they are recycled to G2. Therefore, in embodiments wherethe combined effluents of G1 carry the monomer in excess, the disclosedprocesses remove the excess monomer stream from the effluent of G1 priorto mixing the effluent with the product stream of G2. The removed excessmonomer stream is recycled to their corresponding G1 reactor trainwithout separating and recovering the individual monomer present in itsessentially pure form. The said excess monomer recovery is achieved byphase separation performed in a high-pressure separator as described inU.S. Patent Application No. 60/876,193. It comprises the removal ofenough monomer from the effluent of G1 to ensure that after thereduction, the polymer-containing stream originated from G1 bring in noexcess monomer flow into the recycle streams sent to G2. It means thatafter removing the monomer excess, the polymer-containing product streamoriginated from G1 delivers monomer to the blending point with theeffluent stream of G2 at a rate not higher than the combined purge ratesof the corresponding monomer for the entire process plus the conversionrate of the same monomer in G2. The monomer-rich recycle streamsemerging downstream of the mixing point of G1 and G2 product streams arerecycled to G2. Consequently, at least a part of the monomer content ofthe product stream from G1 will be delivered to G2. Note that since noexcess monomer is brought in by stream mixed with the effluent of G2, noexcess inventory of the monomer builds up in G2 even without recoveringsome or all of the said monomer from the monomer-rich recycle streamsrecovered downstream of the said mixing point between the productstreams of G1 and G2. As mentioned before, the separation processesdisclosed herein achieve the removal of excess monomer from the reactoreffluent of G1 by performing phase separation during which amonomer-rich phase and a reduced polymer-containing (polymer-enriched)phase are recovered from the effluent of G1. Since these phaseseparation is performed on a stream originated exclusively from G1, itwill be referred to as G1 phase separator, or G1 separator throughoutthe current disclosure. Recycling the thus-separated monomer-rich streamto G1 ensures that no excess monomer stream is directed toward G2.Naturally, since no foreign monomer is present in this monomer-rich G1recycle stream, there is no need for monomer separation before itsrecycle to G1. The reduced polymer-enriched stream is mixed with the(unreduced) product stream of G2 to ultimately yield an in-line polymerblend product. The reduction of the effluent originated from G1, i.e.the recovery of excess monomer stream recycled to G1, is controlled byadjusting the temperature and pressure of the phase separator so thatthe combined flow rate of the monomer component derived from G1 andblended with the product streams of G2 is not higher than the combinedpurge rate of the said monomer component for the entire process plus thecombined conversion rate of the said monomer component in G2. Theadjustment of the conditions of the phase separator for recovering theG1 monomer recycle streams can be readily performed by using standardchemical engineering techniques, including phase behavior measurements,well known in the art. The product-containing streams of G1 and G2 arecombined upstream of or in the high-pressure separator downstream of theG1 phase separator as described in U.S. Patent Application No.60/876,193. Since this high-pressure separator has a dual functionseparating a monomer-rich stream from a polymer-rich stream whileblending the monomer and polymer components originated from G1 and G2,it is also referred to as separator-blender. The monomer-rich streamemerging from the separator-blender is recycled to G2. A part of thisand other monomer recycle streams of the process may be purged tocontrol the buildup of undesirable heavy and light components, such asexcess solvents, catalyst killers, light alkanes, etc. The polymer-richstream emerging from the separator-blender is sent to the finishingsection of the in-line blending process disclosed in U.S. PatentApplication No. 60/876,193. The monomer-rich streams emerging downstreamof the separator-blender may be combined with the monomer-rich streamemerging from the separator-blender for recycle to G2. They may also bepartially or fully be purged from the system to control the buildup ofundesirable components. Note that this embodiment corrects for themonomer excess present in the combined effluent streams of G1 and G2 bypartially recovering the excess monomer by phase separation prior toblending the polymer-containing product streams of G1 and G2 in theseparator-blender of the process described in U.S. Patent ApplicationNo. 60/876,193. Note also that the said separator-blender of the processdescribed in U.S. Patent Application No. 60/876,193 is the firsthigh-pressure separator among the serially connected separators thatcombine all the polymer-containing effluents of G1 and G2. It is alsoimportant to emphasize that the monomer-rich effluents originated fromand downstream of the said separator-blender are recycled to G2 and notto G1. The monomer-rich streams recovered by the single-line G1separator upstream of the said separator-blender, on the other hand, isalways recycled directly to its own reactor train belonging to G1 (alsowithout monomer separation). The destination of the recycled monomerstream recovered in a high-pressure separator, therefore, divides thehigh-pressure separators into two groups: one sends back the recoveredmonomer-rich streams to G1, while the other one to G2. The first amongthe serially connected separators of the novel monomer recycle processesdescribed herein that send their monomer-rich recycle streams to G2 isalso the first that combines all polymer-containing product streams.Since it combines the functions of polymer-monomer separation andpolymer-polymer blending, it is also referred to as separator-blender.We need to point out that in our descriptions the separator-blenderdesignation is reserved to this one unique, readily identifiablehigh-pressure separator. Therefore, the earlier-mentioned separatorsending monomer-enriched stream to the reactor train in G1 is nevercalled separator-blender in our descriptions. Instead, its is referredto as G1 phase separator or G1 separator. It is important to point outthat monomer recoveries in the disclosed processes are achieved by phaseseparations as opposed to the conventional cryogenic distillationrequired for the separation of monomer recycle streams emergingdownstream of the point where the product streams of G1 and G2 aremixed. Phase separations are simpler, require lower investment, andoperate with less energy, making the processes disclosed hereinadvantageous over the conventional recycle methods requiring costlydistillative separation and recovery of the individual monomers.

In all embodiments of the novel monomer recycle process operated inconduction with the in-line fluid phase polymer blending processesdescribed in detail in U.S. Patent Application No. 60/876,193, thepolymer-rich stream emerging from the bottom of the above-describedseparator blender is further processed in the finishing section wherethe remaining of the lights are removed to ultimately yield the in-lineproduced polymer blend product. The first stage of the said finishingsection typically comprises one or more phase separators, often referredto as low-pressure separators, from which the volatiles, such asmonomers and solvents, emerge as monomer-rich streams at the top, whilethe polymer-containing stream emerges as further-enriched polymer blendat the bottom. Some or all of the monomer-rich streams emerging from theproduct finishing section, i.e., downstream of the separator-blender,may be combined with the monomer-rich phase of the separator-blender forrecycle to G2 and some or all of these monomer-rich streams may bepurged from the process to control the buildup of volatile inerts, suchas light alkanes, N₂, solvent, etc., in G2. Since the separators in thefinishing section operate at lower pressures than the separator-blender,the monomer-rich recycle streams recovered in the finishing section arecompressed/pumped to higher pressure before merging them with themonomer-rich recycle stream of the separator-blender. The last separatorof the finishing stage is typically followed by a devolatalizer toremove the last traces of the volatile residues to yield the productpolymer blend. The devolatalizer typically operates under vacuum andsometimes also serves as an extruder.

Optionally, in any embodiment of the novel recycle processes operatingin conjunction with the in-line fluid phase polymer blending processesdescribed in detail in U.S. Patent Application No. 60/876,193, theundesired heavy components present in the monomer recycle streams andtypically comprising excess solvents, catalyst killers (e.g., water,methanol, ethanol, etc.), low molecular weight oligomeric side-products,etc. can be separated from one or more monomer recycle streams toprevent the buildup of inert (e.g., solvents, low molecular weightpolymers, etc.) or harmful components (such as excess catalyst killers,etc.) in the reactor trains receiving monomer recycle streams. Theseparation of heavy components is typically simple due to the largedifferences between the boiling points of the light monomers, likeethylene, ethylene, propylene, and butenes and that of the heavycomponents and thus may simply require a knock-out pot or a simpleseparation tower.

Similarly, in order to prevent the buildup of undesired lightcomponents, such as, e.g., inerts originated from the fresh olefinfeeds, an optional light purge stream may be removed from any or allrecycle streams. The one or more light purge streams may advantageouslybe taken from the lowest-pressure monomer recoveries coming from vesselsdownstream of the high-pressure separator-blender described earlier,i.e., from the finishing section of the in-line blending process of U.S.Patent Application No. 60/876,193. Such purge methods save the cost ofrecompression of those lower-pressure monomer-containing recoverystreams. Note that also common to all embodiments of the recycleprocesses disclosed herein is that the blended monomer streams emergingfrom and downstream of the separator-blender (i.e., downstream of thepoint where the polymer-containing effluents originated from G1 and G2are mixed) are combined and recycled to only one of the reactor trains(i.e., to G2) of the in-line blending process of U.S. Patent ApplicationNo. 60/876,193 without separating and recovering the individual monomersfrom the said combined monomer recycle stream.

In all embodiments that apply phase separators to recover monomerrecycle streams for reactor trains in G1, said G1 high-pressureseparators may also serve as buffer vessels to improve the control ofthe polymer blend component ratio. The use and operation of separatorsas buffer vessels is described in detail in U.S. Patent Application No.60/876,193, incorporated herein in its entirety.

In all embodiments of the novel monomer recovery and recycle processesdisclosed herein, the phase separators need to operate in thefluid-fluid two phase regime but need to avoid conditions at whichpolymeric solids would separate. Since prior to the phase separationstep the streams to be separated are in their single fluid phase state,their process conditions need to be changed to bring them to theirfluid-fluid two phase state while avoiding conditions that would lead tothe formation of a solid polymer phase. U.S. Patent Application No.60/876,193 teaches that the desired transition from a single fluid stateto a fluid-fluid two-phase state may be advantageously achieved bypressure letdown. However, in some embodiments, when the temperatures ofthe streams brought to the phase separators are not high enough, suchpressure letdowns may cross the fluid-solid phase transition line in thephase diagrams of the fluid stream leading to the formation of polymersolids and thus ultimately causing fouling. In these embodiments, theheat content of a fluid stream to be separated is increased by passingit through a heat exchanger prior to the pressure letdown.

Two or more polymer-containing product streams of the process may beblended upstream of or in the separators. When product streams arebrought directly to the separators for blending, the separatorsadvantageously have a mixing zone or a mixing chamber to enhanceespecially the blending of the polymeric components of the blendedstreams. When two or more product streams are blended upstream of theseparators, static mixers may be optionally employed between the mixingpoint and the downstream separator to facilitate the blending of thecomponents of the blended streams. Also, when two or more productstreams are blended upstream of the separators, a letdown valve may alsobe deployed downstream of the mixing point, i.e., in the line carryingthe combined streams. The deployment of a pressure letdown valvedownstream of the mixing point may be advantageous for (a) bringing thedifferent streams to a common pressure that is above the cloud point,thus affording the mixing of the phases before fluid-fluid phaseseparation begins, or (b) providing a high-shear mixing of streams thatare already in the fluid-fluid two-phase zone at the point of mixing.When the mixing point is upstream of a letdown valve, there may or maynot be additional letdown valves in some or all stream lines beingdirected to the mixing point. These additional letdown valves may bedeployed to stage the pressure letdown process and/or to bring thestreams to be mixed to a common intermediate pressure that is lower thanthe corresponding upstream (e.g., reactor) pressure but higher than thedownstream separator pressure. Such pressure staging may affordadvantages in heat exchange (by, for example, allowing heat exchange atlower pressure and thus at lower cost) or in mixing efficiency (seeabove).

The monomer recycle process of the present disclosure offers animproved, lower-cost separation-recycle method in which the monomersfrom one or more effluents of the first group of reactor trains (G1) ofthe in-line fluid phase polymer blending process are partially recoveredby phase separation in the form of olefin-enriched streams. The saidpartially recovered monomers are directly recycled to the said firstgroup of reactor trains (G1). The one or more polymer-enriched reducedeffluent streams originated from the first group of reactor trains (G1)are then blended with the other (unreduced) effluent streams of thefirst (G1) and second (G2) groups of reactor trains of the in-line fluidphase polymer blending process. The volume and composition of themonomer blend streams recovered after mixing the effluents of G1 and G2can be adjusted by the degree of (partial) monomer recovery from theeffluents of the first group of reactor trains (G1) and are set suchthat they can be recycled to the second group of reactor trains (G2)without building an excess inventory of the monomers common to both thefirst (G1) and second (G2) group of reactor trains. Thus, the processesof the present disclosure combine the advantages of solution blending ofresin components without the need for costly separation and recovery ofindividual, pure monomer streams before their recycle to thepolymerization reactors. The main advantage of the novel recycleprocesses disclosed herein is that they balance the monomer flows of anin-line polymer blending process and thus enable monomer recycle withoutthe application of costly cryogenic monomer separations. Instead of theexpensive separation of individual monomer streams, the processesdisclosed herein employ phase separators that cost less to install,maintain, and operate.

In essence, the monomer recycle methods for use with fluid phase in-linepolymer blending processes disclosed herein comprise a polymerizationsection and one or more monomer-polymer separator vessels, one of whichcombines all polymer-containing effluent streams of G1 and G2 and servesboth as polymer-polymer blender and polymer-monomer separator, calledthe separator-blending vessel, or separator-blender. The separationvessels sometimes are also referred to high-pressure separators. Asmentioned above, the separator-blender serves as both a separator and asa blender for the reactor effluents of the two or more parallel reactortrains in the reactor bank comprising G1 and G2. Although any and allcombinations of reactor operation modes may be included in the monomerrecycle methods for use with fluid phase in-line polymer blendingprocesses disclosed herein, it is advantageous when at least one of thereactor trains operates in a homogeneous fluid phase and moreadvantageous when all reactors operate in the homogenous fluid phase foreconomic and process simplicity reasons. Bulk homogeneous fluid phasepolymerizations such as bulk homogeneous supercritical or bulk solutionpolymerizations are particularly advantageous.

The monomer recycle methods for use with fluid phase in-line polymerblending disclosed herein offer significant advantages relative to priorart methods of recycling monomers. One or more of the advantages of themonomer recycle process of the present disclosure include, but are notlimited to, reduced capital investment cost by averting the need forrefrigerated monomer separation towers, reduced operating costs byaverting the need for utility requirements associated with operatingrefrigerated monomer separation towers, and reduced maintenance costsdue to process simplicity. In addition, the fluid phase in-line polymerblending process offers one or more of the advantages including, but arenot limited to, improved polymer blend homogeneity because ofmolecular-level mixing of blend components, improved cost of manufacturebecause of savings from avoidance of the reprocessing cost associatedwith conventional off-line blending processes that start with theseparately produced solid, pelletized polymer blend components, andbecause of the ease and simplicity of blending polymers at substantiallyreduced viscosities due to the presence of substantial amounts ofmonomers and optionally solvents in the blending step; flexibility ofadjusting blend ratios and therefore blend properties in-line;flexibility in adjusting production rates of the blend components;flexibility in independently controlling for each reactor the residencetime, monomer composition and conversion, catalyst choice, catalystconcentration, temperature and pressure; improved blend quality;flexibility in making a broader slate of blended products in the sameplant; reduced process cost by utilizing the monomer-polymerseparator(s) for product blending and, in some embodiments, for productbuffering to allow better control of blend ratio.

In-Line Blending Process Overview for Use with Monomer Recycle Process:

Polyolefins are used in a large number of different applications. Eachof these applications requires a different balance between thestiffness, elasticity, and toughness of the polymer. Ideally, polymerswould be custom-tailored to the different needs of each customer. One ofthe methods enabling product tailoring involves the blending ofindividual polymer components. The ability to adjust thestiffness-elasticity-toughness balance of polyolefins provides for theability to meet the needs of a broad range of applications and thus toexpand the potential of polyolefins in delivering desired performance atreduced cost. The stiffness-elasticity-toughness balance may be alteredby changing the molecular structure of polymers by changing theircomposition (i.e. making copolymers), stereoregularity, molecularweight, etc. The stiffness-elasticity-toughness may also be readilyshifted by making blends of polymers or by producing composites. Thein-line blending processes disclosed herein relate to making polymerblends.

Disclosed herein are advantageous monomer recycle processes for use withdirect in-line polymer blend production in an integrated multi-reactorpolymerization wherein the blending step is achieved downstream of thereactors in a separator-blending vessel (also referred to as thehigh-pressure separator). The production of polymer blends in thepolymerization plant is facilitated when the polymer blend componentsare dissolved in the polymerization system since the small-moleculecomponent(s), such as monomer(s) and optional solvent(s)/diluent(s) ofthe polymerization system reduce(s) viscosity thus allowing molecularlevel blending in a low shear process. Hence, using the reactoreffluents wherein the polymer blending components are present in adissolved fluid state may be advantageous to downstream blendingoperations. The polymerization reactors advantageously may be of thehomogeneous supercritical process, the solution process type, or acombination thereof in order to provide the precursor polymer forblending in a fluid state in the direct reactor effluents suitable forthe in-line blending process. Bulk homogeneous supercritical and bulksolution polymerization processes are particularly useful for producingblend components due to the simplicity of the monomer recycle loop anddue to the enhancements in reactor productivity and product properties,such as molecular weight and melting behavior, as will become apparentfrom the following discussions. The in-line blending processes disclosedherein can also utilize certain other polymerization reactors makingin-line blend components, for example, in the form of a slurry, whereinthe polymers form solid pellets in a dense fluid polymerization system.When some of the in-line polymer blending components are made in slurryreactor trains, a dissolution stage is added between the polymerizationreactor train and the separator-blending vessel. This dissolution stagetypically consists of a pump followed by a heater to bring the reactoreffluent above the solid-fluid phase transition conditions affording astream that contains the polymer blending component homogeneouslydissolved in the dense fluid polymerization system. In order tofacilitate the dissolution of the polymer pellets, increased shearingmay be applied, which typically is provided by stirring or by pumping.For embodiments, in which excess monomers need to be taken out from theslurry reactor effluents, the recovery of the desired fraction of themonomer can simply be performed by filtration or by gravimetricseparation of the monomer and the polymer-enriched fraction prior to thepolymer dissolution step. In general, however, because of the addedprocessing and investment costs of such slurry reactor operations,homogeneous polymerization processes, such as homogeneous supercriticalor solution polymerization, are typically cost-advantaged and thusadvantageous to produce the in-line polymer blending components.

As described before, in certain embodiments of the in-line blendingprocesses for use with the monomer recycle process disclosed herein, oneor more reactor effluent streams containing the dissolved polymer blendcomponents are fed to independent separators or separation vessels (alsoreferred to as single-stream high-pressure separators, or Group 1separators, or Group 1 high-pressure separators, or G1 separators, or G1high-pressure separators) upstream of the separator-blending vessel forseparation of a polymer-enriched stream from some fraction of themonomer and the optional solvent/diluent content of the said streams.Such single-stream (G1) high-pressure separators deployed upstream ofthe separator-blending vessel in essence afford a partial recovery ofone or more monomers and the optional solvents present in the effluentsof the first group of reactor trains (G1) thus allowing their recoveryand recycle before being mixed with monomers and optional solvents usedin the reactor trains of the second group of reactor trains (G2). Suchprocesses may be advantageous by eliminating the need for cryogenicallyseparating mixed monomer and optional solvent streams before recyclingthem to the appropriate reactor trains of G2. The polymer-enrichedstreams from each of these single-stream (G1) separators are blended inone of the high-pressure separator vessels that serves both as apolymer-monomer separator and as a polymer-polymer blender(separator-blending vessel). In this embodiment, the operationconditions of the single-stream (G1) separator(s) upstream of theseparator-blending vessel may be adjusted to yield polymer-enrichedstream(s) that still contain(s) enough low molecular weightcomponent(s), such as monomer(s) and optional inert solvent(s) to keepthe viscosity of these streams much below that of the essentially puremolten polymer(s) thus facilitating the mixing of the blending polymercomponents in the separator-blender. Optionally, the G1 separator(s)feeding the separator-blending vessel may also serve as buffer vessel(s)affording an improved control of the blend ratio by compensating for thesmall but inevitable fluctuations in the production of the individualin-line blend components. The buffer capacity of these vessels isdefined by the volume between the operationally allowed maximum andminimum levels of the separated polymer-enriched lower phase.

As opposed to using a cascade of series reactors for the in-lineblending of polymers, the blending processes disclosed herein providefor the individual components of the polymer blend to be made in a bankof parallel reactors. Such direct blend production may be advantageouslyachieved in polymerization processes that operate in a homogeneous densefluid phase, i.e. above the fluid-solid phase transition limits. Theinvention process has at least one reactor train that operates in thehomogeneous dense fluid phase. Polymerization processes that operate ina homogenous dense fluid phase use either inert solvent(s) or monomer(s)or their mixtures as a solvent/diluent in their liquid or supercriticalstate. Hence, such parallel reactors operate with polymerization systemsin their homogeneous dense supercritical or in their liquid state. Inboth the supercritical and liquid operation modes, the process may be abulk polymerization process operating with less than 40 wt %, or lessthan 30 wt %, or less than 20 wt %, or less than 10 wt %, or less than 5wt % of inert solvent present in the feed to the reactor, and in someembodiments, essentially free (less than 1 wt %) of inert solvents. Inone embodiment of the disclosed process, the reactors operate at bulkhomogeneous supercritical conditions as has been disclosed in U.S.patent application Ser. Nos. 11/433,889 and 11/177,004, hereinincorporated by reference in their entirety.

In another embodiment of in-line blending process for use with themonomer recycle process disclosed herein, optionally one or more reactortrains operate at conditions where the polymer dissolution issubstantially aided by an inert solvent (solution process where thepolymerization medium contains more than 40 wt % solvent, typically morethan 60 wt % solvent) as has been disclosed in PCT Publication No. WO2006/044149, herein incorporated by reference in its entirety. In yetanother embodiment, one or more of the reactors included in the bank ofreactors operate in the homogeneous supercritical state and one or moreof the reactors included in the bank of reactors operate in the solutionstate (combination of solution process and homogeneous supercriticalprocess reactors). Both solution and homogeneous supercriticalpolymerization processes provide polymers dissolved in a fluid state,which is required for the downstream in-line blending of polymers. Bothsolution and homogeneous supercritical polymerization processesproviding polymers in a homogeneous fluid state may be performed in abulk monomer phase using essentially pure monomer(s) as solvent or maykeep the polymer in the homogeneous fluid state by employing an inertsolvent in substantial concentrations (i.e., 60 wt % or more). Thesolution process provides for a polymer-containing liquid phase eitherin an inert solvent or in the essentially neat monomer or in theirmixture in their liquid state. The homogeneous supercritical processprovides for the polymeric fluid state by dissolving the polymericproduct either in an inert solvent or in the essentially neat monomer orin their mixture in their supercritical state.

The parallel reactor configuration disclosed herein permits forflexibility in independently controlling for each reactor the residencetime, monomer composition and conversion, catalyst choice, and catalystconcentration not available in a series reactor configuration forblending of polymers. It also makes the independent control of reactiontemperature and pressure easier thus enhancing the control of thepolymerization processes yielding the individual in-line polymer blendcomponents.

U.S. patent application Ser. Nos. 11/433,889 and 11/177,004 disclose aflexible homogeneous polymerization platform for the homogeneoussupercritical propylene polymerization process (also referred to hereinas the “supercritical process”). In the referred supercritical propylenepolymerization process, polymerization is carried out in a substantiallysupercritical monomer medium, thus it is a bulk homogeneoussupercritical polymerization process. The polymer is in a homogeneouslydissolved state in the reactor and in the reactor effluent thus makingthe reactor effluent suitable for a direct downstream blending operationprior to recovering the polymeric products in their solid pelletized orbaled form. U.S. patent application Ser. Nos. 11/433,889 and 11/177,004also teach that the supercritical polymerization process provides anadvantageous means to the so-called solution processes in its ability toproduce highly crystalline, high molecular weight (i.e. low melt-flowrate) isotactic propylene homopolymers. Unlike gas phase and slurrypolymerization processes, the supercritical process may also produceethylene-propylene copolymers and propylene homopolymers with reducedtacticity, and thus reduced polymer melting point without fouling. Aspreviously referenced, U.S. patent application Ser. Nos. 11/433,889 and11/177,004 are incorporated by reference in their entirety herein.

Advantageous polymer blends are often composed of a blend of (a) highlycrystalline component(s) and (a) low crystallinity component(s). Inparticular, solution polymerization processes may provide for lowcrystallinity products because the polymer is present in solution in thereactor, and therefore cannot foul it. However, the solution process haslimitations in producing highly crystalline, high molecular weightproducts with higher melting point. One particularly relevant limitationof the solution process is that it typically cannot produce high MWproducts that also have high melting point, and if it could, suchproducts tend to crystallize in the reactor and cause fouling. Incontrast, the homogeneous supercritical process may provide for bothhigh crystallinity/high melting point and low crystallinity/low meltingpoint polymers without fouling. It also generates the polymer blendcomponents in a dissolved state in the polymerization system allowingdirect blending without the need for a dissolution step. Theseattributes make it a particularly advantageous polymerization processfor the in-line blending processes disclosed herein. Notwithstanding,any combination of polymerization processes operating with densepolymerization systems may be deployed in the in-line blending processesdisclosed herein as long as at least one of the reactor trains operateswith a homogeneous polymerization system. Homogeneous operation isensured by operating above the solid-fluid phase transition point,advantageously not lower than 10 MPa, or not lower than 1 MPa, or notlower than 0.1 MPa, or not lower than 0.01 MPa below the cloud point ofthe polymerization system.

The monomers for use in the bank of parallel reactors disclosed hereinmay be any olefinic compounds containing at least one aliphatic doublebond. The olefin group may be unsubstituted or substituted by one ormore aliphatic or aromatic group(s) and may be part of an open chain ora non-aromatic ring. Exemplary, but not limiting, olefins include alphaand internal linear or branched olefins and their blends, such asethylene, propylene, butenes, pentenes, hexenes, heptenes, octenes,nonenes, decenes, styrenes, non-conjugated dienes, cyclohexene,norbornene, and the like. Exemplary, but not limiting, non-polymerizing(inert) fluid components serving as diluents/solvents include lightparaffinic and aromatic hydrocarbons and their blends, such as butanes,pentanes, hexanes, heptanes, octanes, toluene, xylenes, cyclopentane,cyclohexane, fluorocarbons, hydrofluorocarbons, etc.

The conditions in the polymerization reactors of the aforementionedolefin polymerization process may be established such that the entirereactor content, including the monomer(s), optional non-polymerizingfluid, catalyst system(s), optional scavenger(s) and polymeric products,is in a homogeneous fluid, and advantageously in a single homogeneousfluid state. In certain embodiments, the conditions in the reactors ofthe aforementioned process may be set such that the reactor contents arein their supercritical fluid state, and advantageously in a singlehomogeneous supercritical fluid state.

The upper limit for temperature is determined by the product propertiesthat are strongly influenced by the reaction temperature (for anexample, see FIG. 2). Since often polymers with higher molecular weightsand/or higher melting points are desired, high polymerizationtemperatures (>250° C.) are generally not advantageous. Increasedtemperatures can also degrade most known catalytic systems, providinganother reason for avoiding excessive polymerization temperatures. Atthe current state of the art of polymerization, polymerizationtemperatures above 350° C. are not recommended. For the slurrypolymerization processes, the upper temperature limits of polymerizationare also influenced by the solid-fluid phase transition conditions sincerunning near the solid-fluid phase transition line leads to fouling. Forthat reason, slurry operations not higher than 5° C. below thesolid-fluid phase transition are advantageous, not higher than 10° C.below the solid-fluid phase transition are particularly advantageous.

The lower limits of reaction temperature are determined by the desiredpolymer properties. Lower temperatures generally favor highercrystallinity and higher molecular weight (for an example, see FIG. 2).For homogeneous polymerization processes, the lower limits of reactiontemperature are also determined by the solid-fluid phase transitiontemperature. Running the reactors below the solid-fluid phase transitiontemperature of the reaction mixture may lead to operation problems dueto fouling. For the production of highly crystalline polypropylenes(melting peak temperatures >150° C.) in bulk homogeneous supercriticalpolymerization processes, the minimum operating temperature is about95-100° C. In the production of lower melting copolymers, such asethylene-propylene and ethylene-hexene-1 copolymers, significantly lowerreactor temperatures, e.g., 90° C. or even lower (80 or 70, or 60, or50, or 40° C.), may be readily used without fouling. The application ofcertain inert solvents may further reduce the minimum operationtemperature of the fouling-free operation regime, although, as discussedearlier, the substantial presence of inert solvents also tends to limitthe product molecular weight and often the melting peak temperature. Italso increases production cost due to the need for solvent handling.

The critical temperature and pressure of the polymerization systems aredifferent from the critical values of pure components, and thussupercritical operations at temperatures lower than the criticaltemperature of pure propylene and C4 plus monomers (e.g., 92° C. forpropylene) are possible and disclosed herein. Additionally,near-amorphous and amorphous materials with low melting points may beproduced without fouling even below the critical temperature of thereactor blends, i.e., at temperatures that correspond to the condensedliquid state of the polymerization system in the reactor. In theseinstances, the operation temperature may be below the bubble point ofthe reaction mixture and thus the reactor operates at what is oftenreferred to as liquid-filled conditions. In some instances, suchoperation mode could be desired to achieve high molecular weight (MW)and thus low melt flow rate (MFR), particularly in the manufacture ofcopolymers, such as propylene-ethylene or ethylene-higher olefin orpropylene-higher olefin copolymers. Thus, reactor operations underconditions at which the polymeric products are dissolved in the monomeror monomer blend present in its liquid state, also known as bulksolution polymerization, are also disclosed herein.

Polymerization Temperature for Homogeneous Fluid Phase Polymerizations:

The polymerization process temperature should be above the solid-fluidphase transition temperature of the polymer-containing fluidpolymerization system at the reactor pressure, or at least 2° C. abovethe solid-fluid phase transition temperature of the polymer-containingfluid polymerization system at the reactor pressure, or at least 5° C.above the solid-fluid phase transition temperature of thepolymer-containing fluid polymerization at the reactor pressure, or atleast 10° C. above the solid-fluid phase transformation point of thepolymer-containing fluid polymerization system at the reactor pressure.In another embodiment, the polymerization process temperature should beabove the cloud point of the single-phase fluid polymerization system atthe reactor pressure, or 2° C. or more above the cloud point of thefluid polymerization system at the reactor pressure. In still anotherembodiment, the polymerization process temperature should be between 40and 350° C., or between 50 and 250° C., or between 60 and 250° C., orbetween 70 and 250° C. or between 80 and 250° C. Exemplary lowerpolymerization temperature limits are 40, or 50, or 60, or 70, or 80, or90, or 95, or 100, or 110, or 120° C. Exemplary upper polymerizationtemperature limits are 350, or 250, or 240, or 230, or 220, or 210, or200, or 190, or 180, or 170, or 160° C.

In certain embodiments, polymerization is performed in a supercriticalpolymerization system. In such embodiments, the reaction temperature isabove the critical temperature of the polymerization system. In someembodiments, some or all reactors operate at homogeneous supercriticalpolymerization conditions. Said homogeneous supercriticalpolymerizations of the in-line blending processes disclosed herein maybe carried out at the following temperatures. In one embodiment, thetemperature is above the solid-fluid phase transition temperature of thepolymer-containing fluid reaction medium at the reactor pressure or atleast 5° C. above the solid-fluid phase transition temperature of thepolymer-containing fluid reaction medium at the reactor pressure, or atleast 10° C. above the solid-fluid phase transformation point of thepolymer-containing fluid reaction medium at the reactor pressure. Inanother embodiment, the temperature is above the cloud point temperatureof the single-phase fluid reaction medium at the reactor pressure, or 2°C. or more above the cloud point temperature of the fluid reactionmedium at the reactor pressure. In yet another embodiment, thetemperature is between 50 and 350° C., between 60 and 250° C., between70 and 250° C., or between 80 and 250° C. In one form, the temperatureis above 50, 60, 70, 80, 90, 95, 100, 110, or 120° C. In another form,the temperature is below 350, 250, 240, 230, 220, 210, or 200° C. Inanother form, the cloud point temperature is above the supercriticaltemperature of the polymerization system or between 50 and 350° C.,between 60 and 250° C., between 70 and 250° C., or between 80 and 250°C. In yet another form, the cloud point temperature is above 50, 60, 70,80, 90, 95, 100, 110, or 120° C. In still yet another form, the cloudpoint temperature is below 350, 250, 240, 230, 220, 210, or 200° C.

Polymerization Pressure for Homogeneous Fluid Phase Polymerizations:

The maximum reactor pressure may be determined by process economics,since both the investment and operating expenses increase withincreasing pressure. The minimum pressure limit for the production ofthe individual blend components disclosed herein is set by the desiredproduct properties, such as molecular weight (MW) and melt flow rate(MFR) (see, for example, FIG. 3).

Reducing process pressures in homogeneous polymerizations may lead tophase separation creating a polymer-rich and a polymer-lean fluid phase.In well-stirred reactors, where mass transport is sufficiently high dueto efficient mixing of the two phases, product qualities may not beimpacted by such fluid-fluid phase separation. Therefore, polymerizationprocess conditions under which there is a polymer-rich and apolymer-lean phase are provided herein as long as both phases are abovethe solid-fluid phase separation limit thus preventing fouling and arewell mixed thus preventing substantial mass transfer limitation leadingto poorly controlled increases in molecular weight and/or compositionaldistributions.

Exemplary, but not limiting, process pressures, are between 1 MPa (0.15kpsi) to 1500 MPa (217 kpsi), and more particularly between 1 and 500MPa (0.15 and 72.5 kpsi). Advantageously, the process pressure isbetween 5 and 100 MPa (725 and 14,500 psi) In one embodiment, thepolymerization process pressure should be no lower than the solid-fluidphase transition pressure of the polymer-containing fluid polymerizationsystem at the reactor temperature. In another embodiment, thepolymerization process pressure should be no lower than 10 MPa (or nolower than 1 MPa, or no lower than 0.1 MPa, or no lower than 0.01 MPa)below the cloud point of the fluid polymerization system at the reactortemperature and less than 1500 MPa. In still another embodiment, thepolymerization process pressure should be between 5 and 500 MPa, orbetween 10 and 500 MPa, or between 10 and 300 MPa, or between 20 and 250MPa, or between 10 and 100 MPa. Exemplary lower pressure limits are 1,5, 10, 20, and 30 MPa (0.15, 0.75, 1.45, 2.9, 4.35 kpsi, respectively).Exemplary upper pressure limits are 1500, 1000, 500, 300, 250, 200, and100 MPa (217, 145, 72.5, 43.5, 36.3, 29, and 14.5 kpsi, respectively).

In certain embodiments, polymerization is performed in a supercriticalpolymerization system. In such embodiments, the reaction pressure isabove the critical the pressure of the polymerization system. In someembodiments, some or all reactors operate at homogeneous supercriticalpolymerization conditions. Said homogeneous supercriticalpolymerizations of the in-line blending processes disclosed herein maybe carried out at the following pressures. The supercriticalpolymerization process of the in-line blending processes disclosedherein may be carried out at the following pressures. In one embodiment,the pressure is no lower than the crystallization phase transitionpressure of the polymer-containing fluid reaction medium at the reactortemperature or no lower than 10 MPa below the cloud point of the fluidreaction medium at the reactor temperature. In another embodiment, thepressure is between 10 and 500 MPa, or between 10 and 300 MPa, orbetween 10 and 200 MPa, or between 10 and 100 MPa, or between 20 and 250MPa. In one form, the pressure is above 10, 20, or 30 MPa. In anotherform, the pressure is below 1500, 500, 300, 250, 200, or 100 MPa. Inanother form, the cloud point pressure is between 1 and 500 MPa, orbetween 1 and 300 MPa, or between 1 and 200, or between 1 and 100, orbetween 5 and 250 MPa. In yet another form, the cloud point pressure isabove 1, 5, 10, 20, or 30 MPa. In still yet another form, the cloudpoint pressure is below 1500, 500, 300, 250, 200, or 100 MPa.

Total Monomer Conversion for Homogeneous Fluid Phase Polymerizations:

Increasing the conversion of the total monomer feed in a single-pass inthe individual reactor trains of the parallel reactor bank can reducethe monomer recycle ratio thus can reduce the cost of monomer recycle.Increasing monomer recycle ratios (i.e., the ratio of recycled/totalmonomer feed to the reactor train) require the treatment and recycle oflarger monomer volumes per unit polymer production, which increasesproduction cost. Therefore, higher monomer conversion (lower recycleratios) often provides for improved process economics. However, becausehigh polymer content in the polymerization system, particularly inhomogeneous polymerization systems, yields high viscosities, whichcorrespondingly may make reactor mixing, heat transfer, and downstreamproduct handling difficult, the monomer conversion in a single pass haspractical operation limits. The viscosity of monomer-polymer blends andthus the practical conversion limits can be readily established bystandard engineering methods known in the art (M. Kinzl, G. Luft, R.Horst, B. A. Wolf, J. Rheol. 47 (2003) 869). Single-pass conversionsalso depend on operating conditions and product properties. For example,FIG. 4 shows how increasing conversion reduces the polymer molecularweight. Therefore, monomer conversion may also be constrained by thedesire to increase the molecular weight of the blend component made inthe given reactor train. Exemplary, but not limiting, total monomersingle pass conversions are below 90%, more particularly below 80% andstill more particularly below 60%. Total monomer conversion is definedas the weight of polymer made in a reactor or in a reactor train dividedby the combined weight of monomers and comonomers in the feed to thereactor or reactor train. It should be understood that while high totalmonomer conversion is often limited by product viscosity or by productproperty targets, the conversion of some highly reactive monomercomponents present in some monomer feed blends may be higher than 90%.For example, the single-pass conversion of ethylene inethylene-propylene or in ethylene-higher olefin feed blends may benearly complete (approaching 100%) and is disclosed herein.

As mentioned above, another factor limiting the total monomer conversionis the MW-decreasing effect of conversion (see FIG. 4). Therefore, theproduction of polymer blend components with high MW requires themoderation of monomer conversion in a single pass beyond that of whatviscosity and other practical operation considerations would dictate.Hence, for the production of blend components with high molecular weight(particularly those with higher than >200 kg/mol weight-averagedmolecular weight −M_(w)), the total monomer conversion may need to bebelow 30%. Again, the conversion of some highly reactive components in amonomer feed blend may be higher, and may even approach 100%.

The single-pass conversion in the polymerization reactors disclosedherein may be adjusted by the combination of catalyst concentration andtotal feed flow rate. The total feed rate determines the averageresidence time (in a back-mixed reactor equal to the reactor volumedivided by the total volumetric flow rate of the effluent). The sameconversion may be achieved at lower residence time by increasing thecatalyst concentration in the feed and vice versa. Lower catalystconcentration may reduce catalyst cost, but may also reduce volumetricproductivity thus requiring higher residence times, and ultimately alarger reactor and thus higher investment cost for the same polymerproduction capacity. The optimum balance between residence time/reactorvolumes and catalyst concentration may be determined by standardengineering methods known in the art. A wide-range of polymer blendcomponents may be produced in the reactors disclosed herein at reactorresidence times ranging from 1 sec to 120 min, particularly from 1 secto 60 min, more particularly from 5 sec to 30 min, still moreparticularly from 30 sec to 30 min, and yet still more particularly from1 min to 30 min. In yet another form of the in-line blending processembodiments disclosed herein, the residence time in the reactorsdisclosed herein may be less than 120, or less than 60, or less than 30,or less than 20, or less than 10, or less than 5, or less than 1minute(s).

In certain embodiments, some or all reactor trains of the inventionprocess operate at supercritical conditions advantageously athomogeneous supercritical conditions, or bulk homogeneous supercriticalconditions. The residence times in the supercritical polymerizationreactors, particularly in the bulk homogeneous supercritical reactorsdisclosed herein are generally lower than the residence times insolution, gas phase, and slurry processes due to the high reaction ratesachieved at the conditions of the supercritical polymerization process.In-line blending processes disclosed herein applying bulk homogeneoussupercritical polymerization often choose residence times between 1 and60 min, and more particularly between 1 and 30 min.

The polymerization reactors of the monomer recycle process for fluidphase in-line polymer blending may be grouped into reactor(s) making asingle blending component, called the reactor train. The reactors of theparallel reactor trains producing all the polymer blend components arereferred to as reactor bank. The reactors in the individual trains andin the entire bank can be of any type useful for making polymers (for areview of different polymerization reactors see Reactor Technology by B.L. Tanny in the Encyclopedia of Polymer Sci. and Eng., Vol. 14, H. F.Mark et al., Eds., Wiley, New York, 1988, and J B P Soares, L C Simon inthe Handbook of Polymer Reaction Engineering, T. Meyer and J.Keurentjes, Eds., Wiley-VCH, Weinheim, 2005, p. 365-430.) and can beconstructed the same way or can be different. The optimal reactor typeand configuration can be determined by standard techniques well known inthe art of polymer reactor engineering.

It should be recognized that the catalytic activity and thus thevolumetric productivity in the individual reactors may be different. Ifthe reactor effluents for in-line blending are directly blended, thecatalytic activity and the volumetric productivity may determine thereactor sizes required for the production of the individual polymerblend components. In order to reduce cost, a single plant may need toproduce several polymer blends with different polymer components blendedover a range of blend ratios. Consequently, a parallel reactor bank willoften have reactors of different sizes allowing for a flexible and thusmore cost effective configuration for the production of differentpolymer blend grades. The optimal reactor volumes may be determined fromthe combination of the composition of the target polymer blends and thevolumetric reactor productivity data using optimization methods known inthe art of chemical engineering.

In commercial practice, reactor productivity tends to vary to somedegree, which in turn may lead to the corresponding level of variabilityin polymer blend ratios. In one embodiment, buffer tanks may be added tothe process downstream of the reactors comprising the bank of parallelreactors, but before the polymer mixing or blending point to compensatefor the fluctuations of the volumetric productivity in each reactortrain producing the individual blend components. The buffer tanks mayimprove the compositional control of the final product blends byhomogenizing the individual reactor effluents and by allowing a moreindependent metering of the polymer blend components. When an individualreactor train effluent is stored in the buffer tank in its liquid stateat a pressure below its bubble point, essentially the entire volume ofthe buffer tank is available for compensating for the differences in theblending and production rates. However, when the individual reactoreffluent is stored in the buffer tank in its supercritical state or inits liquid state but at pressures above its bubble point, the denseliquid or supercritical fluid fills the entire tank. In such operationmodes, the buffering capacity, i.e. the capacity to deviate from theinstant reactor flow rate, is more limited and is associated with thepressure/density changes allowed in the buffer tank and with the size ofthe buffer tank. In the latter case, the process streams may be drivenby a gradual pressure drop downstream of the reactor to avoid the costof installing and operating booster pumps. However, booster pumps may bealternatively installed and operated within the process to increase thepressure range and thus the buffering capacity of the system. When nobooster pumps are deployed, the pressure of the buffer tank should belower than that of the reactor, but higher than that of the linesdownstream of the blending point.

Apparently, while feasible, controlling this kind of buffer system isdifficult and it is not very efficient. Thus, in another embodiment,when the individual reactor effluent is stored in the buffer tank in itssupercritical state or in its liquid state but at pressures above itsbubble point, the conditions in the buffer tanks may be set to achievefluid-fluid phase separation (separator-buffer tank operation).Buffering in this mode can be achieved by allowing the fluid level ofthe denser polymer-rich phase to move up and down between the minimumand maximum levels allowed for the desired level of separation whiletaking the monomer-rich upper phase out of the separator buffer via apressure control valve. One skilled in the art can see that thisoperation mode is analogous to the operation of a buffer tank filledwith a liquid phase containing the polymeric product and a vapor phasecontaining the more volatile components, such as monomer(s) andsolvent(s). In the supercritical regime, the upper phase is apolymer-lean supercritical fluid, while the lower phase is apolymer-rich supercritical fluid, the latter of which can be withdrawnfor blending at a controlled rate required for making a constant blendratio, independent of the short-term fluctuations in the productionratios of the individual blend components. A similar analogy may bederived for liquid-filled operations. The polymer content, and thus theviscosity of the polymer-rich phase can be controlled by properlyadjusting the temperature at constant pressure or by adjusting thepressure at constant temperature in the separator-buffer tank(s). Inthis embodiment, the polymer-rich effluent(s) of the separator-buffertank(s) are combined with the direct, unseparated effluent of one of thereactor trains upstream of the separator-blending vessel that recoversthe monomer of the direct reactor effluent as a supernatant and thein-line polymer blend as the bottom phase. In this particularembodiment, one of the separators serves as a separator-blender, whilethe rest of the separators serve as separator-buffers. Advantageously,the separator-blender is connected to a reactor train making a copolymerblending component while the separator-buffer vessels are connected tothe homopolymer trains. This configuration allows the recovery andrecycle of a significant portion of unmixed monomer(s) without the needfor separations from monomers foreign to the correspondingpolymerization train. This form of the fluid-phase in-line blendingprocess for improved monomer recycle will be described in greater detailbelow under the sub-heading “Monomer Recycle Process Configuration forUse with Fluid Phase In-Line Polymer Blending.”

In another embodiment of the processes disclosed herein, polymeradditives may be added to the polymer blend at ratios of up to 40 wt %,or up to 30 wt %, or up to 20 wt %, or up to 10 wt %, or up to 5 wt % tofurther improve product quality and product properties. Exemplary, butnot limiting polymer additives, include specialty polymers includingpolar polymers, waxes, polyalfaolefins, antioxidants, plasticizers,clarifiers, slip agents, flame retardants, heat and uv stabilizers,antiblocking agents, fillers, reinforcing fibers, antistatic agents,lubricating agents, coloring agents, foaming agents, tackifiers,organically modified clays such as are available from Southern Clay, andmaster batches containing above components. Hence, one or more polymeradditive storage tanks containing liquid, molten, or dissolved polymercomponents and polymer additives may be added to the processes disclosedherein. If solvent(s) is used in these polymer additive storage tanks,it should be the same as used in the polymerization reactors previouslydescribed in order to avoid an increase in separation costs in thesolvent recovery and recycle section of the process. For example, whenthe polymer synthesis process is performed in supercritical propylene,the off-line produced polymer additives may also be advantageouslydissolved in supercritical propylene. Solvent-free introduction of thepolymer additive components may be used when the additive component isbrought into its molten state or when the additive component is a liquidat ambient temperatures.

The homogeneous supercritical polymerization and the solutionpolymerization processes are particularly suitable for providing theproduct polymer in a dissolved fluid state. In one particularembodiment, the supercritical polymerization process is performed in thesubstantial absence of an inert solvent/diluent (bulk homogeneoussupercritical polymerization) and provides the product in a dissolvedsupercritical state for the downstream in-line separation-blendingprocess. More particularly, the supercritical polymerization ofpropylene is performed in the substantial absence of an inertsolvent/diluent (bulk homogeneous supercritical propylenepolymerization) and provides the product in a dissolved supercriticalstate for the downstream in-line separation-blending process.

The total amount of inert solvents is generally not more than 80 wt % inthe reactor feeds of the invention process. In some embodiments, wherethe feed essentially comprises the monomer or monomer blend, like forexample, bulk slurry, or bulk supercritical, or bulk solutionpolymerizations, the minimization of solvent use is desired to reducethe cost of monomer recycling. In these cases, the typical solventconcentration in the reactor feed is often below 40 wt %, or below 30 wt%, or below 20 wt %, or below 10 wt %, or below 5 wt % or even below 1wt %. In one form disclosed herein, the polymerization system comprisesless than 40 wt % aromatic hydrocarbons and advantageously less than 40wt % toluene. In another form disclosed herein, the polymerizationsystem comprises less than 40 wt % saturated aliphatic hydrocarbons andadvantageously less than 40 wt % of hexanes, or pentanes, or butanes,and propane, or their mixtures.

Monomer Recycle Process Configuration for Use with Fluid Phase In-LinePolymer Blending:

As previously described in Detailed Description, the improved monomerrecycle process for use with the fluid phase in-line polymer blendingmay have different detailed process configurations. For example, thenumber of parallel reactor trains and their configurations in theparallel reactor bank may be varied. Typically, each reactor trainserves to produce a single in-line blend component. A given train of theparallel reactor bank may be configured as a single reactor, or two ormore reactors in series. From a practical commercial plant designstandpoint, however, there should be a minimum number of reactors for agiven train of the parallel reactor bank in order to make a givenpolymer blend component. Generally, not more than ten series reactorsare utilized and more particularly not more than three series reactorsare generally utilized in a given reactor train. The number of trainscomprising the earlier-defined first (G1) and second (G2) group ofreactor trains may be two, three, four or five or more in each group.The number of reactors in the parallel reactor bank comprising G1 and G2may be any number, although for economic reasons the number of reactorsshould be maintained as low as the desired product grade slate and plantcapacity allows. The optimum number of parallel reactor trains (alsoreferred to as legs or branches of the reactor bank) may be determinedby standard chemical engineering optimization methods well known in theart. Most typically, the polymerization-blending plant will have two orthree parallel polymerization reactor trains or legs or branches in thereactor bank producing product blends with the corresponding number ofin-line polymer blend components. In the embodiments of the improvedrecycle methods disclosed herein, the reactor trains of the in-lineblending process are divided between the earlier-defined first group ofreactor trains (G1) and the earlier-defined second group of reactortrains (G2). At least one of the reactor trains of the in-line blendingprocesses employing the improved recycle methods disclosed hereinbelongs to the said first group of reactor trains (G1) and at least oneother reactor train belongs to the said second group of reactor trains(G2). For example, and as described before, in one exemplary embodiment,the number of parallel trains in the reactor bank is two with oneparallel train producing an olefin homopolymer and the second paralleltrain producing an olefin copolymer wherein the olefin copolymer has onecomonomer in common with the monomer in the olefin homopolymer paralleltrain. Three or more parallel reactors/legs may be employed if theproduction of the target product blends so requires. Besides the in-linepolymer blend components, the final polymer blends often containadditives and modifiers that are not produced within the samepolymerization process. Therefore, it should be understood that thenumber of components in the final product polymer blend typically ishigher than the number of reactor trains or the number of in-linepolymer blend components.

As mentioned above, the fluid phase in-line blending processes for usewith the monomer recycle processes disclosed herein may also optionallyincorporate other polymers, and polymer additives that were producedoutside the reactor bank of the processes disclosed herein. The optionalother polymer and polymer additive components may first be transferredinto solution or molten fluid state before being blended with thein-line produced polymer blend components. These other polymer andpolymer additive components may be stored in polymer additive storagetanks containing liquid, molten, or dissolved polymer components andpolymer additives prior to being transferred and metered to theseparation-blending vessel (if dissolved in a monomer solvent common tothe homopolymerization and copolymerization trains) or to a mixing pointupstream (if again dissolved in a monomer solvent common to thehomopolymerization and copolymerization trains) or downstream of theseparation-blending vessels. Polymer and polymer additive components maybe accurately metered to the blending vessel or to another mixing pointby one or more pumps or if the downstream pressure is lower, through theuse of one or more pressure letdown valves. The optional additives andmodifiers can be mixed into the product upstream of or directly in theseparator-blending vessel or downstream of the separator-blending vesselof the processes disclosed herein. In order to simplify monomertreatment in the monomer recycle train and thus to reduce the cost ofmonomer recycle, it is often advantageous to add the additives andmodifiers downstream of the separator-blending vessel. In suchembodiments, the additives and modifiers may be mixed with the in-lineproduced polymer blend in dedicated pieces of equipment or in thehardware of the product finishing section of the processes disclosedherein, for example, in the devolatizer extruders.

FIG. 5 exemplifies the prior art recycle process with two reactor trains(1 and 2) in which the effluents of the two reactor trains are broughtto the separator-blender for recovering the bulk of the monomers presentin the combined effluents of the two reactor trains. The polymer-richeffluent of the separator-blender containing the in-line polymer blendcomponents is sent to the finishing stage to remove the remainingmonomers and other volatiles before recovering the in-line polymer blendin a form suitable for shipping and sales. The monomer-rich effluent ofthe separator-blender is sent to a cryogenic separation section torecover the individual monomer components before recycling them toReactor trains 1 and 2. The distillation train also employs a tower or aknock-out vessel to separate a heavy purge stream that comprisessolvents, oligomeric side products, excess catalyst killers, etc. Thedistinctive feature of this prior art form is the application of one ormore cryogenic distillation towers to enable the rebalancing of themonomer feeds thus rendering the compositions of the feeds to Reactortrains 1 and 2 suitable for making the desired in-line polymer blendingcomponents. This separation step is costly because it requires one ormore refrigerated (or chilled) towers to fractionate the two or moremonomers in the combined monomer recycle stream based upon differencesin boiling point temperatures (FIG. 5 exemplifies a case for twomonomers separated in Separation tower (2)). It should be understoodthat the exemplary process scheme depicted in FIG. 5 serves only tohighlight the differences between the prior art processes and thus is asimplified conceptual version of the processes encompassing prior artrecycle methods.

FIG. 6 exemplifies the embodiments in which the monomer components inthe effluent of Reactor (1) are all present also in Reactor effluent (2)and the monomer flows in Reactor effluent (1) are not higher than thepurge rates plus the conversion rates of the corresponding monomers inReactor (2). As described earlier in detail, such embodiments allow therecycle of the combined monomer flows of Reactor effluents (1) and (2)to Reactor (2) without building an excess inventory of any of themonomers present in Reactors (1) and (2). The polymer-containingunreduced Reactor effluents (1) and (2) are optionally treated withcatalyst killers to stop the polymerization downstream of the reactor,where the conditions are not the same as those in the reactors and thuswould yield polymer fractions with different polymer properties (c.f.FIGS. 2, 3, and 4). Such polymer fractions may lead to undesired changesin the performance of the final polymer blend product and thus theapplication of catalyst killers may be advantageous. Although in FIG. 6the optional catalyst killers are added to the reactor effluents afterthe pressure letdown valves, they can be added to the reactor effluentsat any point of the process downstream of the reactors. Thus, certainembodiments would add the catalyst killer at the exit of the reactor orsome other parts of the process. Also, catalyst killers may beintroduced at multiple injection ports to improve their efficiency or tocontrol their stoichiometry if so desired. Reactor effluents (1) and (2)are brought to the high-pressure separator, called Separator-blender,optionally via one or more static mixers for polymer-monomer separationand for polymer-polymer blending. Optionally (not shown in FIG. 6), oneor both of the reactor train effluents may be heated before the pressureletdown in order to maintain the temperature in the downstream lines andin the Separator-blender at the desired value, i.e., above thesolid-fluid phase transition temperature of the polymer-rich phase butbelow the cloud point of the streams entering the separators, includingthe separator-blender, to allow the formation of a polymer-rich denserfluid phase and a monomer-rich lighter fluid phase. TheSeparator-blender recovers a monomer-rich phase that is recycled toReactor (2). As depicted in FIG. 6, this monomer recycle stream may becombined with other monomer-rich streams recovered downstream of theSeparator-blender in the product finishing stage. Since the low-pressureseparators downstream of the Separator-blender operate at lower pressurethan the Separator-blender, the streams recovered from those separatorsand in the finishing section are pressurized before they are combinedwith the monomer-rich stream emerging from the Separator-blender. Themonomer recycle streams (or as depicted in FIG. 6, the single combinedmonomer recycle stream) may be optionally passed through a separationtower or knock-out pot that removes a heavy purge stream comprisingundesired heavy components present in the recovered monomer recyclestreams, such as oligomeric heavies, excess solvent and/or catalystkiller, etc. Not shown in FIG. 6, a light purge stream may also beoptionally removed from the monomer recycle stream going to Reactor (2)to control the buildup of light inert components, like methane, ethane,propane, nitrogen, etc., typically introduced with the monomer feeds asimpurities. The optional light purge stream may be advantageously takenfrom the low-pressure monomer recoveries in the finishing section thussaving the compression cost associated with bringing the lower-pressurestreams to the pressure of Monomer blend recycle stream. Before feedingthe optionally purified Monomer blend recycle stream to Reactor (2), itis pumped to the pressure of the fresh reactor feed. Note that Reactor(1) receives only fresh feeds, while Reactor (2) receives fresh makeupmonomer feeds and the Monomer blend recycle stream. Both reactors alsoreceive catalyst feeds.

FIG. 7 exemplifies the same process scenario as described above for theprocess shown in FIG. 6, except that facilities are provided for feedingoptional polymer and other additives to the product stream. FIG. 7 showsan example in which the additives are all introduced at one point,upstream of the optional static mixer. It should be understood, however,that the processes disclosed herein may introduce additives at multiplepoints and may also introduce them at one or more points anywheredownstream of the first pressure letdown valves after the polymerizationreactors. The point of introduction of additives is selected based onthe best engineering compromise between easier mixing in the presence ofviscosity-reducing light components, such as the monomers, and simpler,more robust monomer recycle achieved when the additives are introducedcloser to the final finishing step thus reducing their carryover intothe monomer recycle streams or even avoiding that carryover altogetherby introducing them downstream of the recovery of the last monomerrecycle stream. Naturally, the optimum point will also depend on thenature of the additives. Some, like polymers, would be carried with thepolymer-rich fraction and would not harm the catalyst, thusadvantageously are introduced more upstream. Some, particularlyinhibitors, could harm the catalyst, and thus care must be taken withtheir introduction to avoid negative impact on catalyst activity and theoverall process stability.

FIG. 8 depicts another exemplary embodiment of the monomer recycleprocess for use with fluid phase in-line polymer blending disclosedherein in which there are two parallel Reactor trains (1 and 2). Reactortrain 1, representing G1, has a dedicated separator vessel (Separator(1)) and may be used to produce, for example, an olefin homopolymer (forexample polypropylene, polyethylene or polybutene). It may also producea copolymer, for example ethylene-propylene or ethylene-hexene-1, etc.However, in the latter case, Reactor (2) needs to have at least threemonomers, two of which need to be common with the two used in Reactor(1). Reactor (2), representing G2, has a separator-blender associatedwith it where the polymer-rich phases from the two parallel reactortrains are combined. In this embodiment, Reactor effluent (1) deliversmore monomer than what is purged from the process plus consumed inReactor train (2) of the same monomer. Thus, a (G1) single-streamhigh-pressure separator (Separator (1)) employed to recover the excessmonomer flow as a monomer-rich stream from the Reactor effluent (1) andthus to balance the monomer recycle streams to avoid the buildup ofexcess monomers in Reactor (2). The monomer-rich phase recovered inSeparator (1) is recycled to Reactor (1). In order to control the rateof monomer recovery in Separator (1), it typically operates at asomewhat higher pressure than the one downstream high-pressure separatorthat serves both as a separator and as a blender (Separator-blender).Therefore, there is an optional pressure letdown valve between Separator(1) and the Separator-blender. Catalyst killing agent may be optionallyintroduced prior to or into Separator 1 and/or the Separator-blender tominimize further polymerization outside the polymerization reactors.Optionally, one or more static mixers positioned before theSeparator-blender vessel, but downstream of the mixing point, may beutilized to enhance mixing between the polymer-rich phases of Reactortrains (1) and (2) associated with the Separator-blender. Optionally(not shown in FIG. 8), one or both of the reactor train effluents may beheated before the first pressure letdown in order to maintain thetemperature in the downstream lines and in the separators, including theseparation-blending vessel, at the desired value, i.e., above thesolid-fluid phase transition temperatures of the polymer-rich phases butbelow the cloud point of the streams entering the separators, includingthe separator-blender, to allow the formation of polymer-enriched orpolymer-rich denser fluid phases and monomer-rich lighter fluid phases.The process of this embodiment may be advantageous in, the production ofpolymer blends since the monomer(s) recovered in the separator dedicatedto G1 may be recycled to the corresponding G1 reactor train(s) withoutthe complex separation from other monomers as was associated with singleseparation-blending vessel operation previously described in FIG. 5.Hence, one advantage of this embodiment is that monomer recycle issimplified and thus affords lower cost in the monomer recycle loop.While multiple separation vessel operation increases cost in theseparator section, it adds flexibility and reduces costs in the monomerrecycle loops. As mentioned before, the novel application of G1separators disclosed herein is advantageous because it applieslower-cost monomer recovery by phase separation performed before monomerpools are merged to balance monomer streams in the recycle loop asopposed to the prior art cryogenic separation necessary once the monomerpools merged. It should be understood that for the sake of avoidingunnecessary complexity of the examples provided herein, the processdepicted in FIG. 8 has only one train in both groups (G1 and G2) ofreactor trains. However, the operations for multiple parallel trains ineither or both groups of reactor trains can be readily derived from thescheme shown in FIG. 8.

Referring again to FIG. 8, Separator (1) separates a monomer-rich phasefrom a polymer-rich phase for Reactor train (1). Reactor train (1)typically makes one or more olefin-based homopolymers, although ingeneral, one or more trains of G1 may also produce copolymers with twoor more monomers. In the specific case depicted in FIG. 8, themonomer-rich phase from Separator (1) is recycled directly to Reactortrain (1) without the need for one or more separation towers and theassociated costs and complexity. The polymer-enriched phase fromSeparator (1) is combined with Reactor effluent (2) emerging fromReactor train (2) and then fed to the Separator-blender. Reactor train(2) is typically producing an olefin-based copolymer the monomer pool ofwhich comprises at least all the monomers of Reactor train (1) (but mayhave a larger monomer pool than the monomer pool of G1). In the specificembodiment shown in FIG. 8, the Separator-blender separates amonomer-rich phase containing residual monomer from Reactor train (1)and comonomers from Reactor train (2) from the polymer-rich blend phasecontaining a blend of olefin based homopolymer and olefin basedcopolymer. In cases when Reactor train (1) comprises two or more reactortrains (not shown in FIG. 8) and two or more of them uses the samemonomer pool, the effluents of the trains with the same monomer poolsmay optionally be combined before the said effluents are sent to the G1phase separator for partial monomer recovery and recycle to one or moreof the corresponding trains in G1. It should be noted that as othertrains in G1, the monomer pools of said parallel reactor trains with thesame monomer pool may have not only one, but also more than one members.A special case of the above configuration is when all reactor trains ofG1 have the same monomer pool. In these embodiments the effluents of allG1 reactor trains may optionally be combined before bringing them forpartial monomer recovery for recycle to G1.

As shown in FIG. 8, the monomer-rich phase emerging from theSeparator-blender may be fed to a single separation tower or knock-outpot to remove undesired heavy components, such as excess solvent andcatalyst killer, oligomeric side-products, etc. from the monomer recyclestream at the bottom of the said tower or knock-out pot. The monomerstream emerging from the top of the tower is then recycled back to thecopolymerization Reactor train (2). Optionally (not shown in FIG. 8), alight purge stream can also be removed from the recycle stream beforefeeding it to Reactor (2). Advantageously, light purge streams may alsobe taken from the streams emerging downstream of the Separator-blender.Such purge configuration can reduce the compression cost associated withbringing the monomer recycle streams recovered in the finishing sectionto the pressure of the optional heavies separation tower/Monomer blendrecycle stream. Overall, the embodiment depicted in FIG. 8 may be morecomplicated and higher cost in the separator section relative to FIG. 5,but may be simpler and lower cost in the monomer recycle loops becauseof the decreased need for separation towers.

FIG. 9 depicts an embodiment that is configured and operates as theprocess shown in FIG. 8, except that the separator serving Reactor train(1) serves also as a buffer vessel. The use and operation of separatorsas buffer vessels is described in detail in U.S. Patent Application No.60/876,193, incorporated herein in its entirety.

FIGS. 10 and 11 show specific embodiments that operate as the processesdemonstrated in FIG. 8 or 9, respectively, except for the facilities forstoring and feeding polymeric and other additives in solution or inmolten form for blending with the polymer blend product made by theprocess. The processes in FIGS. 10 and 11 bring in the additivesupstream of the optional static mixer. In practice, additives may beblended in at different points downstream of the reactor trains and maybe mixed in at one or more than one locations. Generally, the earlier,i.e., the further upstream, these additives are mixed with the processstreams, the easier the mixing is due to the lower viscosity of theprocess streams. However, many of the additives may be poisonous to thecatalyst or may cause other complications in the operations of thein-line blending process. For that reason, they may be brought in later,i.e., further downstream, to minimize or even eliminate the possibilityof such negative process impacts. As it will be appreciated by thoseexperienced in the art of polymerization technology, the optimal mixingpoint for additives is specific to the additive-process combinations andcan be determined by applying standard chemical and polymer engineeringmethods.

As will be appreciated by one skilled in the art of chemicalengineering, the process schematic details of the design of the monomerrecycle processes for use with fluid phase in-line polymer blendingdisclosed herein in terms of reactor configuration, separatorconfiguration, valving, heat management, etc. may be set differentlywithout deviating from the spirit of the monomer recycle and in-lineblending processes disclosed herein. The choice between differentembodiments of the processes disclosed herein will be driven by productperformance requirements and process economics, which can be readilydetermined by standard engineering techniques. However, the monomerrecycle processes for use with fluid phase in-line polymer blendingdisclosed herein are advantageous-relative to the prior art by thevirtue of reduced capital costs for monomer separation towers, reducedoperating costs due to reduced utility expenses, and reduced maintenancecosts due to process simplification while still maintaining all thebenefits associated with in-line polymer blending (reduced blending costdue to savings in investment and operation costs, and enablingwell-controlled and cost-effective molecular-level blending to yieldenhanced polymer blend performance).

The processes disclosed herein provide an effective recycle pathway formaking in-line polymer blends made by homogeneous olefin polymerizationprocesses, particularly for bulk homogeneous polymerization processes,such as bulk supercritical and bulk solution olefin polymerization, anexample of which is bulk homogeneous supercritical propylenepolymerization in combination with propylene-ethylene orpropylene-butene-1 or propylene-hexene-1 or propylene-octene-1, etc.copolymerization. As will be discussed in more detail below, theefficient separation of monomer and polymer is achieved byadvantageously utilizing the cloud point pressure and temperaturerelationships for the relevant (polymer/olefinic monomer) or(copolymer/olefinic monomer blend); e.g. (polypropylene/propylenemonomer), (ethylene-propylene copolymer/ethylene-propylene monomerblend), etc. mixtures.

For illustration, cloud point curves are shown in FIGS. 12-21 for threedifferent polypropylene samples having different molecular weights andcrystallinities dissolved in propylene (at 18 wt %). (Achieve 1635 PP isa commercially available metallocene-catalyzed isotactic polypropylenehaving a Melt Flow Rate, MFR, (I₁₀/I₂-ASTM 1238, 230° C., 2.16 kg) of 32dg/min (dg=decigram=0.1 g) available from ExxonMobil Chemical Company,Houston, Tex. ESCORENE PP 4062 is a commercially available isotacticpolypropylene having an MFR of 3.7 dg/min, available from ExxonMobilChemical Company, Houston, Tex. PP 45379 is an isotactic polypropylenehaving an MFR of 300 dg/min produced using a supported metallocene in aslurry polymerization process.

Polymer Blend Formulations and Products:

Many different types of polymer blends may be made by the monomerrecycle processes for use with fluid phase in-line polymer blendingdisclosed herein. A major fraction of a blend is defined as 50% or moreby weight of the blend. A minor fraction of a blend is defined as lessthan 50% by weight of the blend. The monomer recycle processes for usewith fluid phase in-line polymer blending disclosed herein areadvantageous for producing olefin-based polymer blends in which thepolymers can be divided into two groups, P1 and P2, for which thefollowing conditions satisfied:

N(P1+P2)=N(P2) and N(P2)≧N(P1)

where N(P1+P2) is the number of monomers in the combined monomer pool ofP1 and P2 polymer groups, and N(P1) and N(P2) are the number of monomersin the first (P1) and second (P2) group of polymers, respectively.

In particular, the monomer recycle processes for use with fluid phasein-line polymer blending disclosed herein are advantageous for producingolefin based polymer blends of a homopolymer and a copolymer or blendsof a homopolymer and terpolymer or blends of a copolymer and aterpolymer, etc. Still more particularly, the monomer recycle processesfor use with fluid phase in-line polymer blending disclosed herein areadvantageous for producing olefin based polymer blends of a homopolymerand a copolymer wherein the homopolymer and copolymer have a commonmonomer. That is the number of feed monomers of the copolymerizationreactor train is equal to the number of feed monomers of the combinedhomo- and copolymerization reactor trains. Non-limiting exemplaryhomopolymer-copolymer blends include polypropylene (PP)homopolymer—ethylene-propylene (EP) copolymer, PPhomopolymer—propylene-butene-1 (PB) copolymer, polyethylene (PE)—EPcopolymer, polybutene—PB copolymer, polybutene—ethylene-butene-1copolymer, etc.

The monomer recycle processes for use with fluid phase in-line polymerblending disclosed herein are also advantageous for producing olefinbased polymer blends of a homopolymer or a copolymer and a terpolymer.Still more particularly, the monomer recycle processes for use withfluid phase in-line polymer blending disclosed herein are advantageousfor producing olefin-based polymer blends of a homopolymer or copolymerand a terpolymer wherein the homopolymer or copolymer and terpolymerhave one or two common monomers, respectively. That is the number offeed monomers of the terpolymerization reactor train is equal to thenumber of feed monomers of the combined homo- or copolymerization andterpolymerization reactor trains. Non-limiting exemplaryhomopolymer-terpolymer blends include PPhomopolymer—ethylene-propylene-butene-1 (EPB) terpolymer, PE—EPBterpolymer, polybutene—EPB terpolymer, etc. Non-limiting exemplarycopolymer-terpolymer blends include EPcopolymer—ethylene-propylene-butene-1 (EPB) terpolymer, PB copolymer—EPBterpolymer, ethylene-butene-1 (EB)—EPB terpolymer, etc.

The weight fractions of the individual polymer components in the blendsmade by the monomer recycle processes for use with fluid phase in-linepolymer blending disclosed herein may be similar or different. Thepolymer blends disclosed herein may also derive similar improvementsfrom combinations of different materials in similar or differentproportions. One non-limiting example of a useful polymer blend made bythe fluid phase in-line blending process disclosed herein includes amajor fraction of a highly crystalline moderate molecular weight polymerand a minor fraction of a very high molecular weight, elastomericpolymer with low or no inherent crystallinity. Another non-limitingexample of a useful polymer blend made by the monomer recycle processesdisclosed herein includes a major fraction of a soft, tough, low meltingpolymer with a minor fraction of a highly crystalline, high meltingpolymer. Still another non-limiting example of a useful polymer blendmade by the fluid phase in-line blending process disclosed hereinincludes a major fraction of a highly crystalline polymer with a minorfraction of a low or non-crystalline polymer where the low ornon-crystalline polymer is non-elastomeric.

The polymer blends made by the monomer recycle processes for use withfluid phase in-line polymer blending disclosed herein may provide forimproved properties, and hence use in a wide array of applications. Onesuch exemplary, but non-limiting application, is in medical applicationsrequiring improved resistance to sterilizing doses of high-energyradiation. A polymer blend useful for this particular application mayinclude from 75 to 99 wt % moderate molecular weight propylenehomopolymer with 1 to 25 wt % of an ethylene plastomer (EP copolymer).Alternatively, the ethylene plastomer may be replaced by apropylene-ethylene copolymer containing from 8-16 wt % ethylene. Theplastomer or high propylene copolymer component of the blend providessuperior initial ductility as well as retention of ductility andtolerance of the sterilizing radiation to the blend while thehomopolymer component imparts excellent strength, stiffness andresistance to deformation at elevated temperature to the blend. Polymerblends of propylene homopolymer and ethylene plastomer orpropylene-ethylene copolymer are generally clearer or nearly as clear asthe unblended propylene homopolymer component.

Still another exemplary, but non-limiting application of where thepolymer blends made by the monomer recycle processes for use with fluidphase in-line polymer blending disclosed herein find application is indevices and packaging materials requiring good impact resistance, andparticularly in low temperature environments. Polymer blends useful forthis particular application may include from 60 to 99 wt % of a stiffpropylene homopolymer and/or a relatively stiff, low comonomercontaining propylene copolymer and 1-40 wt % of an ethylene plastomer,propylene copolymer containing 5-20 wt % of comonomer, orcomonomer-propylene elastomer (like ethylene-propylene rubber). Inapplications requiring clarity, incorporating into the polymer blend aminor fraction of a highly compatible ethylene plastomer or propylenecopolymer known to have a minimal deleterious effect or even a positiveeffect on the clarity of blends with polypropylene may provide for such.Such plastomers comprise those with a refractive index and viscositysimilar to the polypropylene with which they are to be blended.Compatible propylene copolymers are exemplified by propylene-ethylenecopolymers containing less than 16 wt %, less than 11 wt %, or less than6 wt % ethylene units.

Still yet another exemplary, but non-limiting application of where thepolymer blends made by the monomer recycle processes for use with fluidphase in-line polymer blending disclosed herein find application arethose where materials requiring a combination of stiffness and impactresistance and/or a combination of heat resistance and impactresistance. A polymer blend useful for these applications are similar incomposition to the blends specified for impact resistant devices andpackages. More particularly, polymer blends useful for this particularapplication may include from 60 to 99 wt % of a stiff propylenehomopolymer and/or a relatively stiff, low comonomer containingpropylene copolymer and 1-40 wt % of an ethylene plastomer, propylenecopolymer containing 5-20 wt % of comonomer, or comonomer-propyleneelastomer (like ethylene-propylene rubber). The stiffness and heatresistance may be increased by increasing the homopolymer or stiffcopolymer portion of the polymer blend. Correspondingly, the impactresistance may be improved by increasing the plastomer, propylenecopolymer or ethylene-propylene rubber portion of the blend. The desiredbalance of product attributes may be accomplished by a careful balancingof the two components.

Still yet another exemplary, but non-limiting application of where thepolymer blends made by the monomer recycle processes for use with fluidphase in-line polymer blending disclosed herein find application arethose where a device and/or package must be sterilized by hightemperature and also must be soft and able to withstand impact abuseeven at low temperatures. Polymer blends useful for this particularapplication may include from 75-99 wt % of one or more stiff homopolymerand/or copolymer components and 1-25 wt % of one or more plastomers, lowto no crystallinity propylene copolymers, and ethylene-propylenerubbers. Where increasing softness of packages and device is desired,one may use a greater fraction of the one or more soft components in theblend and smaller fraction of the one or more stiff components in theblend. Polymer blends useful for this particular application may alsoinclude a major fraction of the soft components and minor fraction ofthe stiff components. Hence the range of polymer blends may include 5-90wt % of the stiff polymer component and 10-95 wt % of the soft polymercomponent.

Still yet another exemplary, but non-limiting application of where thepolymer blends made by the monomer recycle processes for use with fluidphase in-line polymer blending disclosed herein find application arefilms which are required to melt and form a seal at relatively lowelevated temperature yet still maintain integrity at much highertemperature. The range of blend compositions previously specified forsoft, elevated temperature resistant devices and/or packages would applyfor this particular type of film application. Similar relationshipsbetween competing properties and the relative usages of the relativecomponents would also apply for this application. More particularly, agreater fraction of the stiff polymer component may increase the sealintegrity at higher temperatures, whereas a greater fraction of the softpolymer component may improve seal formation at lower temperatures andseal strength at normal temperatures.

Other polymeric materials may also be utilized as the soft component forthe applications previously described. For example, propylenehomopolymers and low comonomer copolymers containing relatively largequantities of chain defects such as stereo- and/or regio-defects may beused in the aforementioned blends in place of or along with theplastomers, ethylene-propylene copolymers and other modifiers. However,one distinction between the high defect propylene homopolymers and lowcomonomer copolymers and the previously described plastomers,ethylene-propylene copolymers and other modifiers is that high defectpropylene homo- and copolymers provide relatively less improvement oflow temperature ductility in the blends because their glass transitiontemperatures are not depressed below the norm for propylene polymers inthe absence of significant amounts of ethylene comonomer incorporatedinto their chains.

As will be appreciated by one skilled in the art of polymer engineering,variations to the aforementioned polymer blends and their advantageousapplications may be made without deviating from the spirit of thepolymer blends provided by fluid phase in-line blending processdisclosed herein.

Catalyst System Overview:

The monomer recycle processes for use with fluid phase in-line polymerblending disclosed herein may utilize any number of catalyst systems(also referred to as catalysts) in any of the reactors of thepolymerization reactor section of the process. The monomer recycleprocesses for use with fluid phase in-line polymer blending disclosedherein may also utilize the same or different catalysts or catalystmixtures in the different individual reactors of the reactor bank of thepresent invention. It should be understood that by using differentcatalyst systems we mean that any part of the catalyst system can varyand any combination is allowed. For example, the disclosed process mayuse unsupported catalyst systems in some trains while using supportedcatalyst systems in other trains. In other embodiments, the catalystsystems in some reactor trains may comprise aluminoxane (for example,MAO) activator, while comprising non-coordinating anion activators insome other trains. In another embodiments, the catalyst systems in somereactor trains may comprise Ziegler-Natta catalysts, while the catalystsystems in other reactor trains of the invention process may comprisemetallocenes or nonmetallocene metal-centered, heteroaryl ligandcatalyst compounds (where the metal is chosen from the Group 4, 5, 6,the lanthanide series, or the actinide series of the Periodic Table ofthe Elements) catalysts activated by aluminoxane or non-coordinatinganion activators or any combinations thereof.

As disclosed herein, Ziegler-Natta catalysts are those referred to asfirst, second, third, fourth, and fifth generation catalysts in thePropylene Handbook, E. P. Moore, Jr., Ed., Hanser, N.Y., 1996.Metallocene catalysts in the same reference are described as sixthgeneration catalysts.

As disclosed herein, nonmetallocene metal-centered, heteroaryl ligandcatalyst compounds (where the metal is chosen from the Group 4, 5, 6,the lanthanide series, or the actinide series of the Periodic Table ofthe Elements), are defined as the class of nonmetallocenemetal-centered, heteroaryl ligand catalyst compounds, where the metal ischosen from the Group 4, 5, 6, the lanthanide series, or the actinideseries of the Periodic Table of the Elements. Just as in the case ofmetallocene catalysts, the nonmetallocene metal-centered, heteroarylligand catalyst compounds (where the metal is chosen from the Group 4,5, 6, the lanthanide series, or the actinide series of the PeriodicTable of the Elements) catalysts are typically made fresh by mixing acatalyst precursor compound with one or more activators. Nonmetallocenemetal-centered, heteroaryl ligand catalyst compounds (where the metal ischosen from the Group 4, 5, 6, the lanthanide series, or the actinideseries of the Periodic Table of the Elements) catalysts are described indetail in PCT Patent Publications Nos. WO 02/38628, WO 03/040095 (pages21 to 51), WO 03/040201 (pages 31 to 65), WO 03/040233 (pages 23 to 52),and WO 03/040442 (pages 21 to 54), each of which is herein incorporatedby reference.

While the number of different catalyst systems deployed in the disclosedprocesses can be any number, the use of no more than five differentcatalysts and more particularly, no more than three different catalystsin any given reactor is advantageous for economic reasons. Thedeployment of no more than ten catalysts or the deployment of no morethan six catalysts in the reactor bank of the polymerization process isadvantageous for economic reasons. The one or more catalysts deployed inthe reactors may be homogeneously dissolved in the fluid reaction mediumor may form a heterogeneous solid phase in the reactor. In oneparticular embodiment, the catalyst(s) is (are) homogeneously dissolvedin the fluid reaction medium. When the catalyst is present as a solidphase in the polymerization reactor, it may be supported or unsupported.

The process disclosed herein may use any combination of homogeneous andheterogeneous catalyst systems simultaneously present in one or more ofthe individual reactors of the polymerization reactor section, i.e., anyreactor of the polymerization section of the present invention maycontain one or more homogeneous catalyst systems and one or moreheterogeneous catalyst systems simultaneously. The process disclosedherein may also use any combination of homogeneous and heterogeneouscatalyst systems deployed in the polymerization reactor section. Thesecombinations comprise scenarios when some or all reactors use a singlecatalyst and scenarios when some or all reactors use more than onecatalyst. The one or more catalysts deployed in the process disclosedherein may be supported on particles, which either can be dispersed inthe fluid polymerization medium or may be contained in a stationarycatalyst bed. When the supported catalyst particles are dispersed in thefluid reaction medium, they may be left in the polymeric product or maybe separated from the product prior to its crystallization from thefluid reactor effluent in a separation step that is downstream of thepolymerization reactor section. If the catalyst particles are recovered,they may be either discarded or may be recycled with or withoutregeneration.

The catalyst may also be supported on structured supports, such as forexample, monoliths comprising straight or tortuous channels, reactorwalls, and internal tubing. When the catalysts are supported, operationmay take place on dispersed particles. When the catalyst is supported ondispersed particles, operations may take place without catalyst recoveryi.e., the catalyst is left in the polymeric product. In anotherembodiment, unsupported catalysts may be dissolved in the fluid reactionmedium.

Catalyst systems may be introduced into the reactor by any number ofmethods. For example, the catalyst may be introduced with themonomer-containing feed or separately. Also, the catalyst(s) may beintroduced through one or multiple ports to the reactor. If multipleports are used for introducing the catalyst, those ports may be placedat essentially the same or at different positions along the length ofthe reactor. If multiple ports are used for introducing the catalyst,the composition and the amount of catalyst feed through the individualports may be the same or different. Adjustment in the amounts and typesof catalyst through the different ports enables the modulation ofpolymer properties, such as for example, molecular weight distribution,composition, composition distribution, and crystallinity.

Catalyst Compounds and Mixtures:

The processes described herein may use any polymerization catalystcapable of polymerizing the monomers disclosed herein if that catalystis sufficiently active under the polymerization conditions disclosedherein. Thus, Group-3-10 transition metals may form suitablepolymerization catalysts. A suitable olefin polymerization catalystshould be able to coordinate to, or otherwise associate with, an alkenylunsaturation. Exemplary, but not limiting, polymerization catalystsinclude Ziegler-Natta catalysts, metallocene catalysts, nonmetallocenemetal-centered, heteroaryl ligand catalyst compounds (where the metal ischosen from the Group 4, 5, 6, the lanthanide series, or the actinideseries of the Periodic Table of the Elements) catalysts and latetransition metal catalysts.

Distinction should be made between active catalysts, also referred to ascatalyst systems herein, and catalyst precursor compounds. Catalystsystems are active catalysts comprising one or more catalyst precursorcompounds, one or more catalyst activators, and optionally, one or moresupports. Catalytic activity is often expressed based on theconcentration of the catalyst precursor compounds without implying thatthe active catalyst is the precursor compound alone. It should beunderstood that the catalyst precursor is inactive without beingcontacted or being treated with a proper amount of activator. Similarly,the catalyst activator is inactive without combining it with a properamount of precursor compound. As will become clear from the followingdescription, some activators are very efficient and can be usedstoichiometrically, while some others are used in excess, and insometimes large excess, to achieve high catalytic activity as expressedbased on the concentration of the catalyst precursor compounds. Sincesome of these activators, for example methylaluminoxane (MAO), increasecatalytic activity as expressed based on the concentration of thecatalyst precursor compounds, they are sometimes referred to as“cocatalysts” in the technical literature of polymerization.

When utilizing homogeneous supercritical polymerization conditions formaking one or more of the in-line blending components, homogenouspolymerization catalysts, such as metallocene-based catalysts and othersingle-site homogenous catalyst systems may be advantageous. Forexample, when polymerizing propylene under supercritical conditions,particularly useful metallocene catalysts and non-metallocene catalystsare those disclosed in paragraphs [0081] to [0111] of U.S. PatentApplication No. 10/667,585 and paragraphs [0173] to [0293] of U.S.patent application Ser. No. 11/177,004, the paragraphs of which areherein incorporated by reference.

The processes disclosed herein can employ mixtures of catalyst compoundsto tailor the properties that are desired from the polymer. Mixedcatalyst systems prepared from more than one catalyst precursorcompounds can be employed in the in-line blending processes to alter orselect desired physical or molecular properties. For example, mixedcatalyst systems can control the molecular weight distribution ofisotactic polypropylene when used with the invention processes or forthe invention polymers. In one embodiment of the processes disclosedherein, the polymerization reaction(s) may be conducted with two or morecatalyst precursor compounds at the same time or in series. Inparticular, two different catalyst precursor compounds can be activatedwith the same or different activators and introduced into thepolymerization system at the same or different times. These systems canalso, optionally, be used with diene incorporation to facilitate longchain branching using mixed catalyst systems and high levels of vinylterminated polymers.

As disclosed herein, two or more of the above catalyst precursorcompounds can be used together.

Activators and Activation Methods for Catalyst Compounds:

The catalyst precursor compounds described herein are combined withactivators for use as active catalysts herein.

An activator is defined as any combination of reagents that increasesthe rate at which a metal complex polymerizes unsaturated monomers, suchas olefins. An activator may also affect the molecular weight, degree ofbranching, comonomer content, or other properties of the polymer.

A. Aluminoxane and Aluminum Alkyl Activators:

In one form, one or more aluminoxanes are utilized as an activator inthe in-line blending processes disclosed herein. Alkyl aluminoxanes,sometimes called aluminoxanes in the art, are generally oligomericcompounds containing —Al(R)—O— subunits, where R is an alkyl group.Examples of aluminoxanes include methylaluminoxane (MAO), modifiedmethylaluminoxane (MMAO), ethylaluminoxane and isobutylaluminoxane.Alkylaluminoxanes and modified alkylaluminoxanes are suitable ascatalyst activators, particularly when the abstractable ligand is ahalide. Mixtures of different aluminoxanes and modified aluminoxanes mayalso be used. For further descriptions, see U.S. Pat. Nos. 4,665,208,4,952,540, 5,041,584, 5,091,352, 5,206,199, 5,204,419, 4,874,734,4,924,018, 4,908,463, 4,968,827, 5,329,032, 5,248,801, 5,235,081,5,157,137, 5,103,031 and European and PCT Patent Publication Nos. EP 0561 476 A1, EP 0 279 586 B1, EP 0 516 476 A, EP 0 594 218 A1 and WO94/10180, all of which are herein incorporated by reference in theirentirety.

When the activator is an aluminoxane (modified or unmodified), someembodiments select the maximum amount of activator at a 5000-fold molarexcess Al/M over the catalyst compound (per metal catalytic site). Theminimum activator-to-catalyst-compound is typically a 1:1 molar ratio.

B. Ionizing Activators:

It is contemplated to use an ionizing or stoichiometric activator,neutral or ionic, such as tri(n-butyl)ammonium tetrakis(pentafluorophenyl)boron, a trisperfluorophenyl borone metalloidprecursor or a trisperfluoronaphtyl borone metalloid precursor,polyhalogenated heteroborane anions (PCT patent publication no. WO98/43983), boric acid (U.S. Pat. No. 5,942,459) or combination thereofas an activator herein. Also contemplated for use herein are neutral orionic activators alone or in combination with aluminoxane or modifiedaluminoxane activators.

Examples of neutral stoichiometric activators include tri-substitutedboron, aluminum, gallium and indium or mixtures thereof. The threesubstituent groups are each independently selected from alkyls,alkenyls, halogen, substituted alkyls, aryls, arylhalides, alkoxy andhalides. The three groups are independently selected from halogen, monoor multicyclic (including halosubstituted) aryls, alkyls, and alkenylcompounds and mixtures thereof, preferred are alkenyl groups having 1 to20 carbon atoms, alkyl groups having 1 to 20 carbon atoms, alkoxy groupshaving 1 to 20 carbon atoms and aryl groups having 3 to 20 carbon atoms(including substituted aryls). Alternately, the three groups are alkylshaving 1 to 4 carbon groups, phenyl, napthyl or mixtures thereof.Alternately, the three groups are halogenated, preferably fluorinated,aryl groups. Alternately, the neutral stoichiometric activator istrisperfluorophenyl boron or trisperfluoronapthyl boron.

Ionic stoichiometric activator compounds may contain an active proton,or some other cation associated with, but not coordinated to, or onlyloosely coordinated to, the remaining ion of the ionizing compound. Suchcompounds and the like are described in European patent publication Nos.EP-A-0 570 982, EP-A-0 520 732, EP-A-0 495 375, EP-B1-0 500 944, EP-A-0277 003 and EP-A-0 277 004, and U.S. Pat. Nos. 5,153,157, 5,198,401,5,066,741, 5,206,197, 5,241,025, 5,384,299 and 5,502,124 and U.S. patentapplication Ser. No. 08/285,380, filed Aug. 3, 1994, all of which areherein fully incorporated by reference.

C. Non-Ionizing Activators:

Activators are typically strong Lewis-acids which may play either therole of ionizing or non-ionizing activator. Activators previouslydescribed as ionizing activators may also be used as non-ionizingactivators.

Abstraction of formal neutral ligands may be achieved with Lewis acidsthat display an affinity for the formal neutral ligands. These Lewisacids are typically unsaturated or weakly coordinated. Examples ofnon-ionizing activators include R¹⁰(R¹¹)₃, where R¹⁰ is a group 13element and R¹¹ is a hydrogen, a hydrocarbyl, a substituted hydrocarbyl,or a functional group. Typically, R¹¹ is an arene or a perfluorinatedarene. Non-ionizing activators also include weakly coordinatedtransition metal compounds such as low valent olefin complexes.

Non-limiting examples of non-ionizing activators include BMe₃, BEt₃,B(iBu)₃, BPh₃, B(C₆F₅)₃, AlMe₃, AlEt₃, Al(iBu)₃, AlPh₃, B(C₆F₅)₃,aluminoxane, CuCl, Ni(1,5-cyclooctadiene)₂.

Additional neutral Lewis-acids are known in the art and will be suitablefor abstracting formal neutral ligands. See in particular the reviewarticle by E. Y.-X. Chen and T. J. Marks, “Cocatalysts forMetal-Catalyzed Olefin Polymerization Activators, Activation Processes,and Structure-Activity Relationships”, Chem. Rev., 100, 1391-1434(2000).

Suitable non-ionizing activators include R¹⁰(R¹¹)₃, where R¹⁰ is a group13 element and R¹¹ is a hydrogen, a hydrocarbyl, a substitutedhydrocarbyl, or a functional group. Typically, R¹¹ is an arene or aperfluorinated arene.

Other non-ionizing activators include B(R¹²)₃, where R¹² is an arene ora perfluorinated arene. Alternately, non-ionizing activators includeB(C₆H₅)₃ and B(C₆F₅)₃. Another non-ionizing activator is B(C₆F₅)₃.Alternately, activators include ionizing and non-ionizing activatorsbased on perfluoroaryl borane and perfluoroaryl borates such as PhNMe₂H⁺B(C₆F₅)₄ ⁻, (C₆H₅)₃C⁺ B(C₆F₅)₄ ⁻, and B(C₆F₅)₃.

Additional activators that may be used with the catalyst compoundsdisclosed herein include those described in PCT patent publication No.WO 03/064433A1, which is incorporated by reference herein in itsentirety.

Additional useful activators for use in the processes disclosed hereininclude clays that have been treated with acids (such as H₂SO₄) and thencombined with metal alkyls (such as triethylaluminum) as described inU.S. Pat. No. 6,531,552 and EP Patent No. 1 160 261 A1, which areincorporated by reference herein.

Activators also may be supports and include ion-exchange layeredsilicate having an acid site of at most −8.2 pKa, the amount of the acidsite is equivalent to at least 0.05 mmol/g of 2,6-dimethylpyridineconsumed for neutralization. Non-limiting examples include chemicallytreated smectite group silicates, acid-treated smectite group silicates.Additional examples of the ion-exchange layered silicate include layeredsilicates having a 1:1 type structure or a 2:1 type structure asdescribed in “Clay Minerals (Nendo Kobutsu Gaku)” written by HaruoShiramizu (published by Asakura Shoten in 1995).

Examples of an ion-exchange layered silicate comprising the 1:1 layer asthe main constituting layer include kaolin group silicates such asdickite, nacrite, kaolinite, metahalloysite, halloysite or the like, andserpentine group silicates such as chrysotile, lizaldite, antigorite orthe like. Additional non-limiting examples of the ion-exchange layeredsilicate include ion-exchange layered silicates comprising the 2:2 layeras the main constituting layer include smectite group silicates such asmontmorillonite, beidellite, nontronite, saponite, hectorite,stephensite or the like, vermiculite group silicates such as vermiculiteor the like, mica group silicates such as mica, illite, sericite,glauconite or the like, and attapulgite, sepiolite, palygorskite,bentonite, pyrophyllite, talc, chlorites and the like. The clays arecontacted with an acid, a salt, an alkali, an oxidizing agent, areducing agent or a treating agent containing a compound intercalatablebetween layers of an ion-exchange layered silicate. The intercalationmeans to introduce other material between layers of a layered material,and the material to be introduced is called as a guest compound. Amongthese treatments, acid treatment or salt treatment is particularlyadvantageous. The treated clay may then be contacted with an activatorcompound, such as TEAL, and the catalyst compound to polymerize olefins.

Polar compounds often act as catalyst poisons. Thus, in another form,the polymerization systems comprise less than 5 weight % polar species,or less than 4 weight %, or less than 3 weight %, or less than 2 weight%, or less than 1 weight %, or less than 1000 wt ppm, or less than 750wt ppm, or less than 500 wt ppm, or less than 250 wt ppm, or less than100 wt ppm, or less than 50 wt ppm, or less than 10 wt ppm. Polarspecies include oxygen containing compounds (except for aluminoxanes)such as alcohols, oxygen, ketones, aldehydes, acids, esters and ethers.

In yet another form, the polymerization systems comprise less than 5weight % trimethylaluminum and/or triethylaluminum, or less than 4weight %, or less than 3 weight %, or less than 2 weight %, or less than1 weight %, or less than 1000 ppm, or less than 750 wt ppm, or less than500 wt ppm, or less than 250 wt ppm, or less than 100 wt ppm, or lessthan 50 wt ppm, or less than 10 wt ppm.

In still yet another form, the polymerization systems comprisemethylaluminoxane and less than 5 weight % trimethylaluminum and ortriethylaluminum, or less than 4 weight %, or less than 3 weight %, orless than 2 weight %, or less than 1 weight %, or less than 1000 wt ppm,or less than 750 wt ppm, or less than 500 wt ppm, or less than 250 wtppm, or less than 100 wt ppm, or less than 50 wt ppm, or less than 10 wtppm.

The processes disclosed herein may use finely divided, supportedcatalysts to prepare propylene/1-hexene copolymers with greater than 1.0mole % 1-hexene. In addition to finely divided supports, in-lineblending processes disclosed herein may use fumed silica supports inwhich the support particle diameter may range from 200 angstroms to 1500angstroms, small enough to form a colloid with reaction media.

Catalyst Supports:

In another form, the catalyst compositions of the monomer recycleprocesses for use with fluid phase in-line polymer blending disclosedherein may include a support material or carrier. For example, the oneor more catalyst components and/or one or more activators may bedeposited on, contacted with, vaporized with, bonded to, or incorporatedwithin, adsorbed or absorbed in, or on, one or more supports orcarriers.

The support material may be any of the conventional support materials.In one form, the supported material may be a porous support material,for example, talc, inorganic oxides and inorganic chlorides. Othersupport materials may include resinous support materials such aspolystyrene, functionalized or crosslinked organic supports, such aspolystyrene divinyl benzene polyolefins or polymeric compounds,zeolites, clays, or any other organic or inorganic support material andthe like, or mixtures thereof.

Useful support materials are inorganic oxides that include those Group2, 3, 4, 5, 13 or 14 metal oxides. In one form, the supports includesilica, which may or may not be dehydrated, fumed silica, alumina (PCTpatent publication No. WO 99/60033), silica-alumina and mixturesthereof. Other useful supports include magnesia, titania, zirconia,magnesium chloride (U.S. Pat. No. 5,965,477), montmorillonite (EuropeanPatent No. EP-B1 0 511 665), phyllosilicate, zeolites, talc, clays (U.S.Pat. No. 6,034,187) and the like. Also, combinations of these supportmaterials may be used, for example, silica-chromium, silica-alumina,silica-titania and the like. Additional support materials may includethose porous acrylic polymers described in European Patent No. EP 0 767184 B1, which is incorporated herein by reference. Other supportmaterials include nanocomposites as described in PCT WO 99/47598,aerogels as described in WO 99/48605, spherulites as described in U.S.Pat. No. 5,972,510 and polymeric beads as described in WO 99/50311,which are all herein incorporated by reference.

The support material, for example an inorganic oxide, has a surface areain the range of from about 10 to about 700 m²/g, pore volume in therange of from about 0 to about 4.0 cc/g and average particle size in therange of from about 0.02 to about 50 μm. Alternatively, the surface areaof the support material is in the range of from about 50 to about 500m²/g, pore volume of from about 0 to about 3.5 cc/g and average particlesize of from about 0.02 to about 20 μm. In another form, the surfacearea of the support material is in the range is from about 100 to about400 m²/g, pore volume from about 0 to about 3.0 mL/g and averageparticle size is from about 0.02 to about 10 μm.

Non-porous supports may also be used as supports in the processesdescribed herein. For example, in a one embodiment the nonporous, fumedsilica supports described in U.S. Pat. No. 6,590,055 may be used and isincorporated by reference herein.

Scavengers:

Compounds that destroy impurities are referred to as scavengers by oneskilled in the art of polymerization. Impurities may harm catalysts byreducing their activity. Scavengers may be optionally fed to thereactor(s) of the in-line blending process disclosed herein. Catalyticactivity may be defined many different ways. For example, catalyticactivity can be expressed as turnover frequency, i.e., the number ofmoles of monomers converted to the product in a unit time by one mole ofcatalyst precursor employed in preparing the active catalyst system. Fora given reactor operating at the same residence time, catalytic activitymay also be measured in terms of catalyst productivity, customarilyexpressed as the weight of polymer made by a unit weight of catalystprecursor with or without the weight of the activator.

The scavengers for use in the processes disclosed herein may bedifferent chemical compound(s) from the catalyst activator. Non-limitingexemplary scavengers include diethyl zinc, and alkyl aluminum compounds,such as trimethyl aluminum, triethyl aluminum, tri-isobutyl aluminum,and trioctyl aluminum. The scavenger may also be the same as thecatalyst activator and is generally applied in excess of what is neededto fully activate the catalyst. These scavengers include, but are notlimited to, aluminoxanes, such as methyl aluminoxane. The scavenger mayalso be introduced to the reactor with the monomer feed or with anyother feed stream. In one particular embodiment, the scavenger isintroduced with the monomer-containing feed. The scavenger may behomogeneously dissolved in the polymerization reaction medium or mayform a separate solid phase. In one particular embodiment, scavengersare dissolved in the polymerization medium.

Polymerization Monomers and Comonomers:

The processes disclosed herein may be used to polymerize any monomerhaving one or more (non-conjugated) aliphatic double bond(s) and two ormore carbon atoms. Monomers for use in the in-line blending processinclude ethylene, propylene, C4 and higher α-olefins (non-limitingexamples include butene-1, hexene-1, octene-1, and decene-1);substituted olefins (non-limiting examples include styrene, andvinylcyclohexane); non-conjugated dienes (non-limiting examples includevinylcyclohexene, dicyclopentadiene); α,ω-dienes (non-limiting examplesinclude 1,5-hexadiene, 1,7-octadiene); cycloolefins (non-limitingexamples include cyclopentene, cyclohexene); and norbornene.

The processes disclosed herein may be used to polymerize any unsaturatedmonomer or monomers including C₃ to C₁₀₀ olefins, alternatively C₃ toC₆₀ olefins, alternatively C₃ to C₄₀ olefins, alternatively C₃ to C₂₀olefins, and alternately C₃ to C₁₂ olefins. The processes disclosedherein may also be used to polymerize linear, branched or cyclicalpha-olefins including C₃ to C₁₀₀ alpha-olefins, alternatively C₃ toC₆₀ alpha-olefins, alternately C₃ to C₄₀ alpha-olefins, alternatively C₃to C₂₀ alpha-olefins, and alternatively C₃ to C₁₂ alpha-olefins.Suitable olefin monomers may be one or more of propylene, butene,pentene, hexene, heptene, octene, nonene, decene, dodecene,4-methyl-pentene-1,3-methyl pentene-1,3,5,5-trimethyl-hexene-1, and5-ethylnonene-1.

In another embodiment of the processes disclosed herein, the polymerproduced herein is a copolymer of one or more linear or branched C₃ toC₃₀ prochiral alpha-olefins or C₅ to C₃₀ ring containing olefins orcombinations thereof capable of being polymerized by eitherstereospecific and non-stereospecific catalysts. Prochiral, as usedherein, refers to monomers that favor the formation of isotactic orsyndiotactic polymer when polymerized using stereospecific catalyst(s).

Other monomers for use with the in-line blending process disclosedherein may also include aromatic-group-containing monomers containing upto 30 carbon atoms. Suitable aromatic-group-containing monomers compriseat least one aromatic structure, alternately from one to three, andalternately a phenyl, indenyl, fluorenyl, or naphthyl moiety. Thearomatic-group-containing monomer further comprises at least onepolymerizable double bond such that after polymerization, the aromaticstructure is pendant from the polymer backbone. The aromatic-groupcontaining monomer may further be substituted with one or morehydrocarbyl groups including but not limited to C₁ to C₁₀ alkyl groups.Additionally two adjacent substitutions may be joined to form a ringstructure. Aromatic-group-containing monomers may also contain at leastone aromatic structure appended to a polymerizable olefinic moiety.Non-limiting exemplary aromatic monomers include styrene,alpha-methylstyrene, para-alkylstyrenes, vinyltoluenes,vinylnaphthalene, allyl benzene, and indene, and alternatively styrene,paramethyl styrene, 4-phenyl-1-butene and allyl benzene.

Non aromatic cyclic group containing monomers may also be used in theprocesses disclosed herein. These monomers may include up to 30 carbonatoms. Suitable non-aromatic cyclic group containing monomers may haveat least one polymerizable olefinic group that is either pendant on thecyclic structure or is part of the cyclic structure. The cyclicstructure may also be further substituted by one or more hydrocarbylgroups, for example, but not limited to, C₁ to C₁₀ alkyl groups.Non-limiting exemplary non-aromatic cyclic group containing monomersinclude vinylcyclohexane, vinylcyclohexene, vinylnorbornene, ethylidenenorbornene, cyclopentadiene, cyclopentene, cyclohexene, cyclobutene, andvinyladamantane.

Diolefin monomers may also be used in the processes disclosed herein.These diolefin monomers include any hydrocarbon structure, oralternatively C₄ to C₃₀, having at least two unsaturated bonds, whereinat least two of the unsaturated bonds are readily incorporated into apolymer by either a stereospecific or a non-stereospecific catalyst(s).The diolefin monomers may also be selected from alpha, omega-dienemonomers (i.e. di-vinyl monomers), alternatively linear di-vinylmonomers containing from 4 to 30 carbon atoms. Non-limiting exemplarydienes include butadiene, pentadiene, hexadiene, heptadiene, octadiene,nonadiene, decadiene, undecadiene, dodecadiene, tridecadiene,tetradecadiene, pentadecadiene, hexadecadiene, heptadecadiene,octadecadiene, nonadecadiene, icosadiene, heneicosadiene, docosadiene,tricosadiene, tetracosadiene, pentacosadiene, hexacosadiene,heptacosadiene, octacosadiene, nonacosadiene, triacontadiene,particularly preferred dienes include 1,6-heptadiene, 1,7-octadiene,1,8-nonadiene, 1,9-decadiene, 1,10-undecadiene, 1,11-dodecadiene,1,12-tridecadiene, 1,13-tetradecadiene, and low molecular weightpolybutadienes (M_(w) less than 1000 g/mol). Non-limiting exemplarycyclic dienes include cyclopentadiene, vinylnorbornene, norbornadiene,ethylidene norbornene, divinylbenzene, dicyclopentadiene or higher ringcontaining diolefins with or without substituents at various ringpositions.

Non-limiting examples of polar unsaturated monomers include6-nitro-1-hexene, N-methylallylamine, N-allylcyclopentylamine,N-allyl-hexylamine, methyl vinyl ketone, ethyl vinyl ketone,5-hexen-2-one, 2-acetyl-5-norbornene, 7-synmethoxymethyl-5-norbornen-2-one, acrolein, 2,2-dimethyl-4-pentenal,undecylenic aldehyde, 2,4-dimethyl-2,6-heptadienal, acrylic acid,vinylacetic acid, 4-pentenoic acid, 2,2-dimethyl-4-pentenoic acid,6-heptenoic acid, trans-2,4-pentadienoic acid, 2,6-heptadienoic acid,nona-fluoro-1-hexene, allyl alcohol, 7-octene-1,2-diol,2-methyl-3-buten-1-ol, 5-norbornene-2-carbonitrile,5-norbornene-2-carboxaldehyde, 5-norbornene-2-carboxylic acid,cis-5-norbornene-endo-2,3-dicarboxylic acid,5-norbornene-2,2,-dimethanol, cis-5-norbornene-endo-2,3-dicarboxylicanhydride, 5-norbornene-2-endo-3-endo-dimethanol,5-norbornene-2-endo-3-exo-dimethanol, 5-norbornene-2-methanol,5-norbornene-2-ol, 5-norbornene-2-yl acetate,1-[2-(5-norbornene-2-yl)ethyl]-3,5,7,9,11,13,15-heptacyclopentylpentacyclo[9.5.1.1^(3,9).1^(5,15).1^(7,13)]octasiloxane,2-benzoyl-5-norbornene, allyl 1,1,2,2,-tetrafluoroethyl ether, acroleindimethyl acetal, butadiene monoxide, 1,2-epoxy-7-octene,1,2-epoxy-9-decene, 1,2-epoxy-5-hexene, 2-methyl-2-vinyloxirane, allylglycidyl ether, 2,5-dihydrofuran, 2-cyclopenten-1-one ethylene ketal,allyl disulfide, ethyl acrylate, methyl acrylate.

Polymerizations may be carried out with any suitable feed composition toyield the desired product composition at an economical single-passconversion. Monomer concentrations are generally lower when substantialamounts of inert solvents/diluents are cofed with the monomers andcatalyst. Although inert solvents/diluents may be used if so desired,low solvent/diluent concentration is often advantageous due to reducedsolvent and monomer recovery-recycle cost. In one embodiment, olefinpolymerization is carried out in the presence of less than 60 wt % ofinert solvent/diluent affording olefin concentrations in the combinedfeeds of the individual reactors of 40 wt % or more, or even 55 wt % ormore, and advantageously 75 wt % or more.

In another embodiment, polymerizations yielding the in-line blendcomponents are carried out in bulk monomer phases, i.e., with combinedreactor feeds comprising inert solvent/diluent at less than 40 wt %, orless than 30 wt %, less than 20 wt %, or less than 15 wt %, or less than10 wt %, or less than 5 wt %, or even less than 1 wt %.

In a particular embodiment, ethylene-propylene copolymer blendcomponents are made with essentially diluent-free monomer feedscontaining 1-18 wt % ethylene and 75-99 wt % propylene. In anotherembodiment, ethylene-propylene copolymer blend components are producedwith essentially diluent-free monomer feeds containing 5-30 wt % ofbutene-1, or hexene-1 and 65-95 wt % of propylene or ethylene.

The processes disclosed herein may be used to produce homopolymers orcopolymers. A copolymer refers to a polymer synthesized from two, three,or more different monomer units. Polymers produced by the processesdisclosed herein include homopolymers or copolymers of any of the abovemonomers.

In one embodiment of the processes disclosed herein, the polymer is ahomopolymer of any C₃ to C₁₂ alpha-olefin, or a homopolymer ofpropylene. In another embodiment the polymer is a copolymer comprisingpropylene and ethylene wherein the copolymer comprises less than 70weight % ethylene, or less than 60 weight % ethylene, or less than 40weight % ethylene, or less than 20 weight 0% ethylene. In anotherembodiment the polymer is a copolymer comprising propylene and one ormore of any of the monomers listed above. In another embodiment, thecopolymers comprise one or more diolefin comonomers, alternatively oneor more C₆ to C₄₀ non-conjugated diolefins, alternatively one or more C₆to C₄₀ α,ω-dienes.

In another embodiment of the processes disclosed herein, the one or morepolymer blend components are a copolymer of ethylene, propylene, orother higher olefin and optionally any third monomer, typically anotherhigher olefin, such as C₄ to C₂₀ linear, branched or cyclic monomers. Inanother embodiment, the one or more polymer blend components producedherein are a copolymer of ethylene and one or more of propylene, butene,pentene, hexene, heptene, octene, nonene, decene, dodecene,4-methyl-pentene-1,3-methyl pentene-1, and 3,5,5-trimethyl hexene 1. Instill another embodiment, the one or more polymer blend componentsproduced herein are a copolymer of propylene and one or more ofethylene, butene, pentene, hexene, heptene, octene, nonene, decene,dodecene, 4-methyl-pentene-1,3-methyl-pentene-1, and3,5,5-trimethyl-hexene-1. In still yet another embodiment, the one ormore polymer blend components produced herein are a copolymer of a C4 orhigher olefin and one or more of ethylene, propylene, butene, pentene,hexene, heptene, octene, nonene, decene, dodecene,4-methyl-pentene-1,3-methyl-pentene-1, and 3,5,5-trimethyl-hexene-1.

In another embodiment of the processes disclosed herein, the copolymersdescribed comprise at least 50 mole % of a first monomer and up to 50mole % of other monomers. In another embodiment, the polymer comprises:a first monomer present at from 40 to 95 mole %, or 50 to 90 mole %, or60 to 80 mole %, and a comonomer present at from 5 to 40 mole %, or 10to 60 mole %, or 20 to 40 mole %, and a termonomer present at from 0 to10 mole %, or from 0.5 to 5 mole %, or from 1 to 3 mole %. Suchcopolymer blending components can be readily produced when thecomonomer(s) is (are) present between 0.1 and 85 mole % in the combinedfeeds to the reactor making the copolymers.

In another embodiment of the processes disclosed herein, the firstmonomer comprises one or more of any C₃ to C₈ linear branched or cyclicalpha-olefins, including propylene, butene, (and all isomers thereof),pentene (and all isomers thereof), hexene (and all isomers thereof),heptene (and all isomers thereof), and octene (and all isomers thereof).Preferred monomers include propylene, 1-butene, 1-hexene, 1-octene,cyclohexene, cyclooctene, hexadiene, cyclohexadiene and the like.

In another embodiment of the processes disclosed herein, the comonomercomprises one or more of any C₂ to C₄₀ linear, branched or cyclicalpha-olefins (provided ethylene, if present, is present at 5 mole % orless), including ethylene, propylene, butene, pentene, hexene, heptene,and octene, nonene, decene, undecene, dodecene, hexadecene, butadiene,hexadiene, heptadiene, pentadiene, octadiene, nonadiene, decadiene,dodecadiene, styrene,3,5,5-trimethylhexene-1,3-methylpentene-1,4-methylpentene-1,cyclopentadiene, and cyclohexene.

In another embodiment of the processes disclosed herein, the termonomercomprises one or more of any C₂ to C₄₀ linear, branched or cyclicalpha-olefins, (provided ethylene, if present, is present at 5 mole % orless), including ethylene, propylene, butene, pentene, hexene, heptene,and octene, nonene, decene, un-decene, dodecene, hexadecene, butadiene,hexadiene, heptadiene, pentadiene, octadiene, nonadiene, decadiene,dodecadiene, styrene,3,5,5-trimethyl-hexene-1,3-methylpentene-1,4-methylpentene-1,cyclopentadiene, and cyclohexene.

In another embodiment of the processes disclosed herein, the polymersdescribed above further comprise one or more dienes at up to 10 weight%, or at 0.00001 to 1.0 weight %, or at 0.002 to 0.5 weight %, or at0.003 to 0.2 weight %, based upon the total weight of the composition.In some embodiments 500 ppm or less of diene is added to the combinedfeed of one or more polymerization trains, alternately 400 ppm or less,alternately 300 ppm or less. In other embodiments at least 50 ppm ofdiene is added to the combined feed of one or more polymerizationtrains, or 100 ppm or more, or 150 ppm or more. In yet anotherembodiment the concentration of diene in the combined feed to thereactor is between 50 wt ppm and 10,000 wt ppm.

In another embodiment of the processes disclosed herein, the processesused to produce propylene copolymers with other monomer units, such asethylene, other α-olefin, α-olefinic diolefin, or non-conjugateddiolefin monomers, for example C₄-C₂₀ olefins, C₄-C₂₀ diolefins, C₄-C₂₀cyclic olefins, C₈-C₂₀ styrenic olefins. Other unsaturated monomersbesides those specifically described above may be copolymerized usingthe processes disclosed herein, for example, styrene, alkyl-substitutedstyrene, ethylidene norbornene, norbornadiene, dicyclopentadiene,vinylcyclohexane, vinylcyclohexene, acrylates, and otherolefinically-unsaturated monomers, including other cyclic olefins, suchas cyclopentene, norbornene, and alkyl-substituted norbornenes.Copolymerization can also incorporate α-olefinic macromonomers producedin-situ or added from another source. Some embodiments limit thecopolymerization of α-olefinic macromonomers to macromonomers with 2000or less mer units. U.S. Pat. No. 6,300,451 discloses many usefulcomonomers. That disclosure refers to comonomers as “a second monomer.”

In another embodiment of the processes disclosed herein, when propylenecopolymers are desired, the following monomers can be copolymerized withpropylene: ethylene, but-1-ene, hex-1-ene, 4-methylpent-1-ene,dicyclopentadiene, norbornene, C₄-C₂₀₀₀, C₄-C₂₀₀, or C₄-C₄₀ linear orbranched, α,ω-dienes; C₄-C₂₀₀₀, C₄-C₂₀₀, or C₄-C₄₀ cyclic olefins; andC₄-C₂₀₀₀, C₄-C₂₀₀, or C₄-C₄₀ linear or branched α-olefins.

Other Primary Monomer:

The polymerization processes disclosed herein may polymerize butene-1(T_(c)=146.5° C.; P_(c)=3.56 MPa), pentene-1 (T_(c)=191.8° C.;P_(c)=3.56 MPa), hex-1-ene (T_(c)=230.8° C.; P_(c)=3.21 MPa), and3-methyl-butene-1 (T_(c)=179.7° C.; P_(c)=3.53 MPa) using these monomersor mixtures comprising the monomers at supercritical conditions or as aliquid. These processes may employ at least one of butene-1, pentene-1,or 3-methyl-butene-1 as monomer. These processes may also employreaction media that comprise butene-1, pentene-1, or 3-methyl-butene-1.These processes can employ polymerization feeds that contain greaterthan 50 mole % of butene-1, pentene-1, or 3-methyl-butene-1 and theirconcentration can vary between 0.1 and 85 mole %. Of course, thesecompounds can be freely mixed with each other and with propylene asmonomer, bulk reaction media, or both.

Polymerization Solvents/Diluents:

Solvent and or diluent may be present in the polymerization system. Anyhydrocarbon, fluorocarbon, or fluorohydrocarbon inert solvent or theirmixtures may be used at concentrations not more than 80 wt % in thefeeds to any individual polymerization reactor of the in-line blendingprocess disclosed herein. The concentration of the inert solvent in thereactor feed and thus in the polymerization system in certainembodiments utilizing bulk polymerization processes is not more than 40wt %, alternatively not more than 30 wt %, alternatively not more than20 wt %, alternately not more than 10 wt %, alternately not more than 5wt %, and alternately not more than 1 wt %.

Diluents for use in the in-line blending process disclosed herein mayinclude one or more of C₂-C₂₄ alkanes, such as ethane, propane,n-butane, i-butane, n-pentane, i-pentane, n-hexane, mixed hexanes,cyclopentane, cyclohexane, etc., single-ring aromatics, such as tolueneand xylenes. In some embodiments, the diluent comprises one or more ofethane, propane, butane, isobutane, pentane, isopentane, and hexanes. Inother embodiments, the diluent is recyclable.

Other diluents may also include C₄ to C₁₅₀ isoparaffins, or C₄ to C₁₀₀isoparaffins, or C₄ to C₂₅ isoparaffins, or C₄ to C₂₀ isoparaffins. Byisoparaffin is meant that the paraffin chains possess C₁ to C₁₀ alkylbranching along at least a portion of each paraffin chain. Moreparticularly, the isoparaffins are saturated aliphatic hydrocarbonswhose molecules have at least one carbon atom bonded to at least threeother carbon atoms or at least one side chain (i.e., a molecule havingone or more tertiary or quaternary carbon atoms), and advantageouslywherein the total number of carbon atoms per molecule is in the rangebetween 6 to 50, and between 10 and 24 in another embodiment, and from10 to 15 in yet another embodiment. Various isomers of each carbonnumber will typically be present. The isoparaffins may also includecycloparaffins with branched side chains, generally as a minor componentof the isoparaffin. The density (ASTM 4052, 15.6/15.6° C.) of theseisoparaffins may range from 0.70 to 0.83 g/cm³; the pour point is −40°C. or less, alternately −50° C. or less, the viscosity (ASTM 445, 25°C.) is from 0.5 to 20 cSt at 25° C.; and the average molecular weightsin the range of 100 to 300 g/mol. Some suitable isoparaffins arecommercially available under the tradename ISOPAR (ExxonMobil ChemicalCompany, Houston Tex.), and are described in, for example in U.S. Pat.Nos. 6,197,285, 3,818,105 and 3,439,088, and sold commercially as ISOPARseries of isoparaffins. Other suitable isoparaffins are also commercialavailable under the trade names SHELLSOL (by Shell), SOLTROL (by ChevronPhillips) and SASOL (by Sasol Limited). SHELLSOL is a product of theRoyal Dutch/Shell Group of Companies, for example Shellsol™ (boilingpoint=215-260° C.). SOLTROL is a product of Chevron Phillips ChemicalCo. LP, for example SOLTROL 220 (boiling point=233-280° C.). SASOL is aproduct of Sasol Limited (Johannesburg, South Africa), for example SASOLLPA-210, SASOL-47 (boiling point 238-274° C.).

In another embodiment of the monomer recycle processes for use withfluid phase in-line polymer blending disclosed herein, diluents mayinclude C₄ to C₂₅ n-paraffins, or C₄ to C₂₀ n-paraffins, or C₄ to C₁₅n-paraffins having less than 0.1%, or less than 0.01% aromatics. Somesuitable n-paraffins are commercially available under the tradenameNORPAR (ExxonMobil Chemical Company, Houston Tex.), and are soldcommercially as NORPAR series of n-paraffins. In another embodiment,diluents may include dearomaticized aliphatic hydrocarbon comprising amixture of normal paraffins, isoparaffins and cycloparaffins. Typicallythey are a mixture of C₄ to C₂₅ normal paraffins, isoparaffins andcycloparaffins, or C₅ to C₁₈, or C₅ to C₁₂. They contain very low levelsof aromatic hydrocarbons, or less than 0.1, or less than 0.01 aromatics.Suitable dearomatized aliphatic hydrocarbons are commercially availableunder the tradename EXXSOL (ExxonMobil Chemical Company, Houston Tex.),and are sold commercially as EXXSOL series of dearomaticized aliphatichydrocarbons.

In another embodiment of the monomer recycle processes for use withfluid phase in-line polymer blending disclosed herein, the inert diluentcomprises up to 20 weight % of oligomers of C₆ to C₁₄ olefins and/oroligomers of linear olefins having 6 to 14 carbon atoms, or 8 to 12carbon atoms, or 10 carbon atoms having a Kinematic viscosity of 10 ormore (as measured by ASTM D 445); and having a viscosity index (“VI”),as determined by ASTM D-2270 of 100 or more.

In another embodiment of the monomer recycle processes for use withfluid phase in-line polymer blending disclosed herein, the inert diluentcomprises up to 20 weight % of oligomers of C₂₀ to C₁₅₀₀ paraffins,alternately C₄₀ to C₁₀₀₀ paraffins, alternately C₅₀ to C₇₅₀ paraffins,alternately C₅₀ to C₅₀₀ paraffins. In another embodiment of the fluidphase in-line process for blending disclosed herein, the diluentcomprises up to 20 weight % of oligomers of 1-pentene, 1-hexene,1-heptene, 1-octene, 1-nonene, 1-decene, 1-undecene and 1-dodecene. Sucholigomers are commercially available as SHF and SuperSyn PAO's(ExxonMobil Chemical Company, Houston Tex.). Other useful oligomersinclude those sold under the tradenames Synfluid™ available fromChevronPhillips Chemical Co. in Pasedena Tex., Durasyn™ available fromBP Amoco Chemicals in London England, Nexbase™ available from Fortum Oiland Gas in Finland, Synton™ available from Chemtura Corporation inMiddlebury Conn., USA, EMERY™ available from Cognis Corporation in Ohio,USA.

In another embodiment of the monomer recycle processes for use withfluid phase in-line polymer blending disclosed herein, the diluentcomprises a fluorinated hydrocarbon. Exemplary fluorocarbons includeperfluorocarbons (“PFC” or “PFC's”) and or hydrofluorocarbons (“HFC” or“HFC's”), collectively referred to as “fluorinated hydrocarbons” or“fluorocarbons” (“FC” or “FC's”). Fluorocarbons are defined to becompounds consisting essentially of at least one carbon atom and atleast one fluorine atom, and optionally hydrogen atom(s). Aperfluorocarbon is a compound consisting essentially of carbon atom andfluorine atom, and includes for example linear branched or cyclic, C₁ toC₄₀ perfluoroalkanes. A hydrofluorocarbon is a compound consistingessentially of carbon, fluorine and hydrogen. FC's include thoserepresented by the formula: C_(x)H_(y)F_(z) wherein x is an integer from1 to 40, alternatively from 1 to 30, alternatively from 1 to 20,alternatively from 1 to 10, alternatively from 1 to 6, alternativelyfrom 2 to 20 alternatively from 3 to 10, alternatively from 3 to 6,alternatively from 1 to 3, wherein y is an integer greater than or equalto 0 and z is an integer and at least one, alternatively, y and z areintegers and at least one. For purposes of the in-line blendingprocesses disclosed herein and the claims thereto, the termshydrofluorocarbon and fluorocarbon do not include chlorofluorocarbons.

In one embodiment of the monomer recycle processes for use with fluidphase in-line polymer blending disclosed herein, a mixture offluorocarbons are used, alternatively a mixture of a perfluorinatedhydrocarbon and a hydrofluorocarbon, and alternatively a mixture of ahydrofluorocarbons. In yet another embodiment, the hydrofluorocarbon isbalanced or unbalanced in the number of fluorine atoms in the HFC used.

In another embodiment of the monomer recycle processes for use withfluid phase in-line polymer blending disclosed herein, the fluorocarbonis not a perfluorinated C4 to C10 alkane. In another embodiment, thefluorocarbon is not perfluorodecalin, perfluoroheptane, perfluorohexane,perfluoromethylcyclohexane, perfluorooctane,perfluoro-1,3-dimethylcyclohexane, perfluorononane, or perfluorotoluene.In another embodiment, the fluorocarbon is present at more than 1 weight%, based upon the weight of the fluorocarbon and any hydrocarbon solventpresent in the reactor, alternately greater than 3 weight %, alternatelygreater than 5 weight %, alternately greater than 7 weight %,alternately greater than 10 weight %, alternately greater than 15 weight%. In some embodiments, the fluorocarbons are present in thepolymerization reaction media at 0 to 20 volume %, based upon the volumeof the media, alternately the fluorocarbons are present at 0 to 10volume %, alternately 0 to 5 volume %, and alternately 0 to 1 volume %.

With regard to the polymerization media, suitable diluents and solventsare those that are soluble in and inert to the monomer and any otherpolymerization components at the polymerization temperatures andpressures.

Polymerization Reactor Configuration:

The polymerization processes of the monomer recycle processes for usewith fluid phase in-line polymer blending disclosed herein may becarried out in two or more parallel reactor trains making the polymersfor downstream blending, and more advantageously in two parallel reactortrains The parallel reactor trains may be fed with essentially the sameor different feeds, although there must be one monomer in common betweenthe two or more parallel reactor trains. The parallel reactor trains maybe run at essentially the same or different reactor conditions. Theparallel reactor trains may also produce essentially the same ordifferent polymeric products, although for downstream blending purposes,the production of different polymeric products is advantageous. Forexample, one parallel reactor train may produce an olefin basedhomopolymer and a second parallel reactor train may produce an olefinbased copolymer (referred to as polymer produced from two, three or morecomonomers), wherein there is one common monomer between the twoparallel reactor trains. The number of feed monomers of thecopolymerization reactor train may be equal to the number of feedmonomers of the combined homo- and copolymerization parallel reactortrains.

When multiple parallel reactor trains are used in the processesdisclosed herein, the production of polymer blends is possible. In oneparticular embodiment, homopolymer and copolymer blends with a commonmonomer are made by using at least two reactor trains in a parallelconfiguration, and advantageously two parallel reactor trains.Non-limiting exemplary homopolymers include polyethylene, polypropylene,polybutene, polyhexene, polyoctene, polydecene, and polystyrene. In oneembodiment, the homopolymer comprises polyethylene, polypropylene,polybutylene, polyhexene, polyoctane, polydecene, and polystyrene. Inanother embodiment, the homopolymer is polyethylene or polypropylene.The copolymers may be any two- or three-component combinations ofethylene, propylene, butene-1, hexene-1, octene-1, decene-1, styrene,norbornene, 1,5-hexadiene, and 1,7-octadiene. In one embodiment, thecopolymers are made from a two-component combination of ethylene,propylene, butene-1, hexene-1, octane-1, decene-1, styrene, norbornene,1,5-hexadiene, and 1,7-octadiene. In another embodiment, the copolymeris an ethylene-propylene, propylene-butene-1, propylene-hexene-1,propylene-octene-1, propylene-decene-1, ethylene-butene-1,ethylene-hexene-1 ethylene-octene-1, and/or ethylene-decene-1 copolymer.In yet another embodiment, the copolymer is anethylene-propylene-butene-1, ethylene-propylene-hexene-1,ethylene-propylene-octene-1, and/or ethylene-propylene-decene-1copolymer.

As previously described, the in-line blending polymer components areproduced in a reactor bank composed of at least two parallel reactortrains and advantageously two parallel reactor trains. A reactor trainof the parallel reactor bank may include one or more reactors that maybe configured in series configuration. The number of parallel reactorstrains or branches in a parallel bank may be any number, but forpractical reasons, is generally limited to less than ten, alternativelynot more than six parallel reactor trains, alternatively not more thanfive or not more than four reactor trains, alternatively not more thanthree parallel reactor trains, and alternatively not more than twoparallel reactor trains. The number of series cascade reactorsconstituting a given reactor train or branch of a parallel configurationmay be any number, but for practical reasons, is generally limited tonot more than ten reactors in series, alternatively not more than sixreactors in series, alternatively not more than three reactors inseries, and alternatively not more than two reactors in series.

In one embodiment, the polymer-containing effluents from two or morereactor trains configured in a parallel configuration are combinedyielding a polymer blend comprising the polymeric products of theindividual reactors without first recovering the polymeric products ofthe individual reactors in solid forms. The two or more reactor trainsconstituting the parallel configuration generally include a singlereactor, or alternatively, two or more reactors in series.

In another embodiment, the polymer-containing effluents from two reactortrains configured in a parallel configuration are combined yielding apolymer blend comprising the polymeric products of the individualreactors without first recovering the polymeric products of theindividual reactors in solid form. The two reactor trains constitutingthe parallel configuration generally include a single reactor, oralternatively, two or more reactors in series.

The reactors of the polymerization system for the monomer recycleprocesses for use with fluid phase in-line polymer blending disclosedherein may be stirred or unstirred. When a reactor train comprises twoor more reactors, the members of the reactor train are not necessarilyconstructed the same way, for example, the individual members of areactor train may be stirred, unstirred, or a combination thereof. Theindividual reactors may also be of equal or different size. The same istrue for the reactors in the entire reactor bank. The optimal reactorconfiguration and sizes may be determined by standard engineeringtechniques known to those skilled in the art of chemical engineering.

Any type of polymerization reactor may be deployed in the monomerrecycle processes for use with fluid phase in-line polymer blendingdisclosed herein. The optimal reactor design may be determined bystandard engineering techniques known to those skilled in the art ofchemical engineering. Non-limiting exemplary reactor designs includestirred tank with or without an external loop, tubular reactor, and loopreactor. The reactors may operate adiabatically or may be cooled. Thecooling may be achieved within the reactor, or through the reactorjacket, or dedicated heat exchange loops may be applied.

Polymerization Process Details:

The monomer recycle processes for use with fluid phase in-line polymerblending disclosed herein relate to processes to polymerize olefinscomprising contacting one or more olefins having at least two carbonatoms with a suitable catalyst compound and an activator in a fluidreaction medium comprising one or two fluid phases in a reactor. In oneembodiment, the fluid reaction medium is in its supercritical state.Catalyst compound and activator may be delivered as a solution orslurry, either separately to the reactor, mixed in-line just prior tothe reactor, or mixed and pumped as an activated solution or slurry tothe reactor. In one particular embodiment, two solutions are mixedin-line. For a given reactor train of the parallel configuration,polymerizations may be carried out in either single reactor operation,in which monomer, comonomers, catalyst(s)/activator(s), scavenger(s),and optional solvent(s) are added continuously to a single reactor or inseries reactor operation, in which the above components are added to twoor more reactors connected in series. The catalyst components may beadded to the first reactor in the series. The catalyst component mayalso be added to each reactor in the series reactor train. The freshcatalyst feed if added to more than one reactor in the series train maybe the same or different to each reactor and their feed rates may be thesame or different.

Polymerization processes of the fluid phase in-line polymer blendingprocesses disclosed herein also comprehend high-pressure reactors wherethe reactor is substantially unreactive with the polymerization reactioncomponents and is able to withstand the high pressures and temperaturesthat occur during the polymerization reaction. Withstanding these highpressures and temperatures may allow the reactor to maintain the fluidreaction medium in its supercritical condition. Suitable reaction vesseldesigns include those necessary to maintain supercritical or otherhigh-pressure ethylene polymerization reactions. Non-limiting exemplaryreactors include autoclave, pump-around loop or autoclave, tubular, andautoclave/tubular reactors.

The polymerization processes of the fluid phase in-line polymer blendingprocesses disclosed herein may operate efficiently in autoclave (alsoreferred to as stirred tank) and tubular reactors. Autoclave reactorsmay be operated in either a batch or continuous mode, although thecontinuous mode is advantageous. Tubular reactors always operate incontinuous mode. Typically, autoclave reactors have length-to-diameterratios of 1:1 to 20:1 and are fitted with a high-speed (up to 2000 RPM)multiblade stirrer and baffles arranged for optimal mixing. Commercialautoclave pressures are typically greater than 5 MPa with a maximum oftypically less than 260 MPa. The maximum pressure of commercialautoclaves, however, may become higher with advances in mechanical andmaterial science technologies.

When the autoclave has a low length-to-diameter ratio (such as less thanfour), the feed streams may be injected at one position along the lengthof the reactor. Reactors with large diameters may have multipleinjection ports at nearly the same or different positions along thelength of the reactor. When they are positioned at the same length ofthe reactor, the injection ports are radially distributed to allow forfaster intermixing of the feed components with the reactor content. Inthe case of stirred tank reactors, the separate introduction of thecatalyst and monomer(s) may be advantageous in preventing the possibleformation of hot spots in the unstirred feed zone between the mixingpoint and the stirred zone of the reactor. Injections at two or morepositions along the length of the reactor is also possible and may beadvantageous. In one exemplary embodiment, in reactors where thelength-to-diameter ratio is from 4 to 20, the reactor may contain up tosix different injection positions along the reactor length with multipleports at some or each of the lengths.

Additionally, in the larger autoclaves, one or more lateral mixingdevices may support the high-speed stirrer. These mixing devices canalso divide the autoclave into two or more zones. Mixing blades on thestirrer may differ from zone to zone to allow for a different degree ofplug flow and back mixing, largely independently, in the separate zones.Two or more autoclaves with one or more zones may connect in a seriesreactor cascade to increase residence time or to tailor polymerstructure in a reactor train producing a polymer blending component. Aspreviously described, a series reactor cascade or configuration consistsof two or more reactors connected in series, in which the effluent of atleast one upstream reactor is fed to the next reactor downstream in thecascade. Besides the effluent of the upstream reactor(s), the feed ofany reactor in the series reactor cascade of a reactor train can beaugmented with any combination of additional monomer, catalyst, orsolvent fresh or recycled feed streams. Therefore, it should beunderstood that the polymer blending component leaving a reactor trainof the process disclosed herein may itself be a blend of the samepolymer with increased molecular weight and/or compositional dispersionor even a blend of homo- and copolymers.

Tubular reactors may also be used in the fluid phase in-line polymerblending processes disclosed herein and more particularly tubularreactors capable of operating up to about 350 MPa. Tubular reactors arefitted with external cooling and one or more injection points along the(tubular) reaction zone. As in autoclaves, these injection points serveas entry points for monomers (such as propylene), one or more comonomer,catalyst, or mixtures of these. In tubular reactors, external coolingoften allows for increased monomer conversion relative to an autoclave,where the low surface-to-volume ratio hinders any significant heatremoval. Tubular reactors have a special outlet valve that can send apressure shockwave backward along the tube. The shockwave helps dislodgeany polymer residue that has formed on reactor walls during operation.Alternatively, tubular reactors may be fabricated with smooth,unpolished internal surfaces to address wall deposits. Tubular reactorsgenerally may operate at pressures of up to 360 MPa, may have lengths of100-2000 meters or 100-4000 meters, and may have internal diameters ofless than 12.5 cm. Typically, tubular reactors have length-to-diameterratios of 10:1 to 50,000:1 and include up to 10 different injectionpositions along its length.

Reactor trains that pair autoclaves with tubular reactors are alsocontemplated within the scope of the fluid phase in-line polymerblending processes disclosed herein. In this reactor system, theautoclave typically precedes the tubular reactor or the two types ofreactors form separate trains of a parallel reactor configuration. Suchreactor systems may have injection of additional catalyst and/or feedcomponents at several points in the autoclave, and more particularlyalong the tube length. In both autoclaves and tubular reactors, atinjection, feeds are typically cooled to near ambient temperature orbelow to provide maximum cooling and thus maximum polymer productionwithin the limits of maximum operating temperature. In autoclaveoperation, a preheater may operate at startup, but not after thereaction reaches steady state if the first mixing zone has someback-mixing characteristics. In tubular reactors, the first section ofdouble-jacketed tubing may be heated (especially at start ups) ratherthan cooled and may operate continuously. A well-designed tubularreactor is characterized by plug flow wherein plug flow refers to a flowpattern with minimal radial flow rate differences. In both multizoneautoclaves and tubular reactors, catalyst can not only be injected atthe inlet, but also optionally at one or more points along the reactor.The catalyst feeds injected at the inlet and other injection points canbe the same or different in terms of content, density, andconcentration. Catalyst feed selection allows polymer design tailoringwithin a given reactor or reactor train and/or maintaining the desiredproductivity profile along the reactor length.

At the reactor outlet valve, the pressure drops to begin the separationof polymer and unreacted monomer, co-monomers, solvents and inerts, suchas for example ethane, propane, hexane, and toluene. More particularly,at the reactor outlet valve, the pressure drops to levels below thatwhich critical phase separation allowing for a polymer-rich phase and apolymer-lean phase in the downstream separation vessel. Typically,conditions remain above the polymer product's crystallizationtemperature. The autoclave or tubular reactor effluent may bedepressurized on entering the downstream high-pressure separator (HPS oralso referred to as a separator, separator vessel, separation vessel,separator/blender vessel, or separation/blending vessel).

As will be subsequently described in detail, the temperature in theseparation vessel is maintained above the solid-fluid phase separationtemperature, but the pressure may be below the critical point. Thepressure need only be high enough such that the monomer may condenseupon contacting standard cooling water. The liquid recycle stream maythen be recycled to the reactor with a liquid pumping system instead ofthe hyper-compressors required for polyethylene units. The relativelylow pressure in separator reduces the monomer concentration in theliquid polymer phase which results in a lower polymerization rate. Thepolymerization rate may be low enough to operate the system withoutadding a catalyst poison or “killer”. If a catalyst killer is required(e.g., to prevent reactions in the high pressure recycle) then provisionmust be made to remove any potential catalyst poisons from the recycledpolymer rich monomer stream for example, by the use of fixed bedadsorbents or by scavenging with an aluminum alkyl.

In an alternative embodiment, the HPS may be operated over the criticalpressure of the monomer or monomer blend but within the densefluid-fluid two phase region, which may be advantageous if the polymeris to be produced with a revamped high-pressure polyethylene (HPPE)plant. The recycled HPS overhead is cooled and dewaxed before beingreturned to the suction of the secondary compressor, which is typical ofHPPE plant operation. The polymer from this intermediate orhigh-pressure vessel then passes through another pressure reduction stepto a low pressure separator. The temperature of this vessel ismaintained above the polymer melting point so that the polymer from thisvessel can be fed as a liquid directly to an extruder or static mixer.The pressure in this vessel is kept low by using a compressor to recoverthe unreacted monomers, etc. to the condenser and pumping systemreferenced above.

In addition to autoclave reactors, tubular reactors, or a combination ofthese reactors, loop-type reactors may be utilized in the fluid phasein-line polymer blending processes disclosed herein. In this reactortype, monomer enters and polymer exits continuously at different pointsalong the loop, while an in-line pump continuously circulates thecontents (reaction liquid). The feed/product takeoff rates control thetotal average residence time. A cooling jacket removes reaction heatfrom the loop. Typically feed inlet temperatures are near to or belowambient temperatures to provide cooling to the exothermic reaction inthe reactor operating above the crystallization temperature of thepolymer product. The loop reactor may have a diameter of 41 to 61 cm anda length of 100 to 200 meters and may operate at pressures of 25 to 30MPa. In addition, an in-line pump may continuously circulate thepolymerization system through the loop reactor.

The polymerization processes of the fluid phase in-line polymer blendingprocesses disclosed herein may have residence times in the reactors asshort as 0.5 seconds and as long as several hours, alternatively from 1sec to 120 min, alternatively from 1 second to 60 minutes, alternativelyfrom 5 seconds to 30 minutes, alternatively from 30 seconds to 30minutes, alternatively from 1 minute to 60 minutes, and alternativelyfrom 1 minute to 30 minutes. More particularly, the residence time maybe selected from 10, or 30, or 45, or 50, seconds, or 1, or 5, or 10, or15, or 20, or 25, or 30 or 60 or 120 minutes. Maximum residence timesmay be selected from 1, or 5, or 10, or 15, or 30, or 45, or 60, or 120minutes.

The monomer-to-polymer conversion rate (also referred to as theconversion rate) is calculated by dividing the total quantity of polymerthat is collected during the reaction time by the amount of monomeradded to the reaction. Lower conversions may be advantageous to limitviscosity although increase the cost of monomer recycle. The optimumtotal monomer conversion thus will depend on reactor design, productslate, process configuration, etc., and can be determined by standardengineering techniques. Total monomer conversion during a single passthrough any individual reactor of the fluid phase in-line process forblending disclosed herein may be up to 90%, or below 80%, or below 60%or 3-80%, or 5-80%, or 10-80%, or 15-80%, or 20-80%, or 25-60%, or3-60%, or 5-60%, or 10-60%, or 15-60%, or 20-60%, or 10-50%, or 5-40%,or 10-40%, or 40-50%, or 15-40%, or 20-40%, or 30-40% or greater than5%, or greater than 10%. In one embodiment, when the product isisotactic polypropylene and long-chain branching (LCB) of thepolypropylene is desired (g′≦0.97 based on GPC-3D and using an isotacticpolypropylene standard), single pass conversions may be above 30% andalternatively single-pass conversions may be above 40%. In anotherembodiment, when isotactic polypropylene essentially free of LCB isdesired (0.97<g′<1.05), single-pass conversions may be not higher than30% and alternatively single-pass-conversions may be not higher than25%. To limit the cost of monomer separation and recycling, single-passconversions may be above 3%, or above 5%, or above 10%. It should beunderstood that the above exemplary conversion values reflect totalmonomer conversion, i.e., the conversion obtained by dividing thecombined conversion rate of all monomers by the total monomer feed rate.When monomer blends are used, the conversion of the more reactivemonomer component(s) will always be higher than that of the lessreactive monomer(s). Therefore, the conversion of the more reactivemonomer component(s) can be substantially higher than the totalconversion values given above, and can be essentially complete,approaching 100%.

Product Separation and Downstream Processing:

The reactor effluents of the processes disclosed herein aredepressurized to a pressure significantly below the cloud pointpressure. This allows separation of a polymer-rich phase for furtherpurification and a monomer-rich phase for optional separation andrecycle compression back to the reactor(s). The reactor effluents may beoptionally heated before pressure let down to avoid the separation of asolid polymer phase, which causes fouling of the separators andassociated reduced-pressure lines. The separation of the polymer-richphase and the monomer-rich phase in the processes disclosed herein iscarried out in a vessel known as a high-pressure separator (alsoreferred to as an HPS, separator, separator vessel, or separationvessel). The high-pressure separator located after the mixing point ofthe polymer-containing product streams of all reactor trains of theparallel reactor bank is also referred to as, separator-blender,separator-blender vessel, or separation-blending vessel recognizing itsdual function of blending the said polymer-containing product streamswhile also separating a monomer-rich phase from a polymer-rich phase,the latter of which comprises the polymer blend of the in-line blendingprocesses disclosed herein.

One or more single-stream high-pressure separators in series areemployed to partially recover the monomer(s) and optional solvent(s)from the effluent of a single reactor train (homopolymerization reactortrain) before sending the polymer-enriched stream to the downstreamseparator-blender. The monomer-rich phase emerging from the top of thehigh-pressure separator may then be directly recycled back to thehomopolymerization reactor train without the need for one or morechilled towers. The downstream separator-blender blends one or morepolymer-enriched streams with one or more unreduced reactor traineffluents (from copolymerization reactor train) to yield a monomer-richphase and a polymer-rich phase, the latter of which comprises thepolymer blend. In one form, the monomer-rich phase may then be directlyrecycled back to the copolymerization reactor train without the need forone or more chilled towers. In another form, the monomer-rich phase maypass through one chilled separation tower if there is a need to separatesolvent, oligomers and other heavies from the monomers prior to recycle.In another embodiment, the single-stream high-pressure separator placedupstream of the separator-blender also functions as a buffer vessel(separator-buffer vessel) by allowing the fluid level of thepolymer-enriched phase to vary in the separator-buffer vessel. Suchbuffering enables a more precise control of the blend ratios bycompensating for the momentary fluctuations in the production rates inthe individual parallel reactor trains of the monomer recycle processesfor use with fluid phase in-line polymer blending disclosed herein.

The polymer-rich phase of the separator-blender may then be optionallytransferred to one or more low-pressure separators (LPS also referred toas a low-pressure separation vessel) running at just above atmosphericpressure for a simple flash of light components, reactants and oligomersthereof, for producing a low volatile-containing polymer melt enteringthe finishing extruder or optional static mixer. The one or morelow-pressure separators are distinguished from the one or morehigh-pressure separators in generally operating at lower pressuresrelative to the high-pressure separators. The one or more low-pressureseparators also are located downstream of the one or more high-pressureseparators including the separator-blender. Moreover, the one or morelow-pressure separators may function to separate light from heavycomponents comprising the polymer blend of the monomer recycle processesfor use with fluid phase in-line polymer blending disclosed herein,whereas the one or more high-pressure separators may function toseparate light from heavy components upstream of the low-pressureseparator (i.e. monomer-rich phase from polymer-rich phase) and mayfunction to blend the polymer-rich phases from two or more parallelreactor trains or may function as buffers. As previously stated, ahigh-pressure separator may be alternatively referred to herein as anHPS, separator, separator vessel, separation vessel, separator-blendervessel, or separation-blending vessel, or separator-blender. The use ofthe term “pressure” in conjunction with low-pressure separator andhigh-pressure separator is not meant to identify the absolute pressurelevels at which these separators operate at, but is merely intended togiven the relative difference in pressure at which these separatorsoperate. Generally, separators located downstream in the monomer recycleprocesses for use with fluid phase in-line polymer blending disclosedherein operate at lower pressure relative to separators locatedupstream.

The monomer recycle processes for use with fluid phase in-line polymerblending disclosed herein for supercritical propylene polymerization maybe carried out under agitation in the single-phase region in the reactorat 40-200 MPa and 95-180° C. (see FIG. 22). The product mixture from oneor more reactor trains (typically the homopolymerization reactor train)may be discharged into a high-pressure separator vessel, where thepressure is dropped to a level of 25 MPa or lower, in which case, themixture is below its cloud point, while the monomer has not yet flashedoff (again, see FIG. 22). Under such conditions, it would be expectedfrom Radosz et al., Ind. Eng. Chem. Res. 1997, 36, 5520-5525 and Loos etal., Fluid Phase Equil. 158-160, 1999, 835-846 that the monomer-richphase would comprise less than about 0.1 wt % of low molecular weightpolymer and have a density of approximately 0.3-0.6 g/mL (see FIG. 23).The polymer-rich phase would be expected to have a density ofapproximately 0.5-0.8 g/mL.

Assuming that the pressure is dropped rapidly enough, for example,greater than or equal to about 6 MPa/sec, the phases will separaterapidly, permitting the recycle of the monomer-rich phase as a liquid,without the issue of having the monomer-rich phase return to the gasphase. As may be appreciated by those skilled in the art, thiseliminates the need for the energy-intensive compression andcondensation steps.

The polymer-rich phase may then be sent to a downstreamseparator-blender for blending with the reactor effluent from thecopolymerization reactor train. The separator-blender separates amonomer-rich phase from a polymer-rich blend phase. The polymer-richblend phase may then be sent to one or more separator in series or maybe sent directly to a coupled devolatilizer. Suitable devolatilizers maybe obtained, for example, from LIST USA Inc., of Charlotte, N.C. Thedevolatilization is a separation process to separate remaining volatilesfrom the final polymer, eliminating the need for steam stripping.Working under low vacuum, the polymer solution flashes into thedevolatilizer, exits the unit and is then transferred on for furtherprocessing, such as pelletization.

Any low or very low molecular weight polymer present in the monomer-richphase to be recycled may optionally be removed through “knock-out” pots,standard hardware in reactors systems, or left in the return streamdepending upon product requirements and the steady-state concentrationof the low molecular weight polymer fraction in the product.

In solution reactor processes, present practices employed by thoseskilled in the art typically effect separation by flashing monomer andsolvent or accessing the high-temperature cloud point.

In another form, polymerization is conducted at conditions below thecloud point, with the polymer-monomer mixture transported to agravimetric separation vessel, where the pressure could be furtherlowered if desired to enhance phase separation of the polymer-rich andmonomer-rich phases. In either of the forms herein described, themonomer, for example, propylene, is recycled while staying in arelatively high density, liquid-like (supercritical or bulk liquid)state. Once again, one or more knock-out pots may be employed to aid inthe removal of low molecular weight polymer from the recycle stream.

As may be appreciated, there are possible and optimal operating regimesfor reactors and for the gravity (lower critical solution temperature(LCST)) separator. Referring now to FIG. 24, for reactors operating in asingle liquid phase regime, a possible region for operation is justabove the LCST and vapor pressure (VP) curves. The optimal region (shownwithin the shaded oval) for operation occurs at temperatures just abovethe lower critical end point (LCEP) and at pressures slightly above theLCST curve.

Referring now to FIG. 25, for reactors operating within a two-phasefluid-fluid regime, the possible region for operation occurs basicallyanywhere below the LCST curve. The optimal region (again, shown withinthe shaded oval) occurs just below the LCST and above the VP curve,although, as may be appreciated, many factors could have a bearing onwhat actually is optimal, such as the final properties of the desiredproduct. As recognized by those skilled in the art, the two-phaseliquid-liquid regime is the economically advantageous method ifpolypropylene is to be produced with a revamped HPPE plant.

Referring now to FIG. 26, for the case where polymerization is conductedat conditions below the cloud point and the polymer-monomer mixturetransported to a gravimetric LCST separator, the possible region ofoperation is anywhere below the LCST curve and above the VP curve. Theoptimal region (again, shown within the shaded oval) occurs within thatportion that is below the spinodal, but not too low in pressure, asshown. Operating in this regime assures that the energy use isoptimized. It is also desirable to avoid operation in the region betweenthe LCST and spinodal curves in order to obtain good gravity settlingperformance. Moreover, it is desirable that the separation is effectedat sufficiently high temperatures, so that crystallization does notoccur in the polymer-rich phase. This may require that the temperatureof the mixture in the separator be higher than the temperature in thereactor(s).

Advantageously, the liquid monomer-rich recycle stream can be recycledto the reactor using a liquid pumping system instead of ahyper-compressor, required for conventional polyethylene units.

Catalyst Killing:

The use of the processes disclosed herein and the relatively lowpressure in the separator vessel greatly reduces the monomerconcentration in the liquid polymer-rich phase, which, in turn, resultsin a much lower polymerization rate. This polymerization rate may be lowenough to operate this system without adding a catalyst poison or“killer”. If no killing compounds are added then the killer removal stepmay be eliminated.

If a catalyst killer is required, then provision must be made to removeany potential catalyst poisons from the recycled monomer-rich stream(e.g. by the use of fixed bed adsorbents or by scavenging with analuminum alkyl). The catalyst activity may be killed by addition of apolar species. Non-limiting exemplary catalyst killing agents includewater, alcohols (such as methanol and ethanol), sodium/calcium stearate,CO, and combinations thereof. The choice and quantity of killing agentwill depend on the requirements for clean up of the recycle propyleneand comonomers as well as the product properties, if the killing agenthas low volatility. The catalyst killing agent may be introduced intothe reactor effluent stream after the pressure letdown valve, but beforethe HPS. The choice and quantity of killing agent may depend on therequirements for clean up of the recycle propylene and comonomers aswell as the product properties, if the killing agent has low volatility.

Polymer Blending Components:

The polymers produced by processes disclosed herein may be manystructure types, including, but not limited to, block, linear, radial,star, branched, and combinations thereof.

Some forms produce polypropylene and copolymers of polypropylene with aunique microstructure. The processes disclosed herein can be practicedsuch that novel isotactic, atactic, and syndiotactic compositions aremade. In other forms, crystalline polymers are made.

The processes disclosed herein produce propylene polymers with a meltingpoint of 40 to 165° C., and a weight-average molecular weight of 2,000to 1,000,000, 10,000 to 1,000,000, 15,000 to 500,000, 25,000 to 250,000or 35,000 to 150,000.

The processes disclosed herein produce polymer with a heat of fusion,ΔH_(f), of 1-30 J/g, 2-20 J/g, or 3-10 J/g. In another form, theprocesses disclosed herein produce polymers having a ΔH_(f) of up to 10J/g, alternately 50 to 110 J/g, alternatively 70 to 10 J/g.

Dienes can be used as a comonomer to increase the molecular weight ofthe resulting polymer and to create long chain branching. Vinyl chloridecan be used as a comonomer to increase the degree of vinyl terminationin the polymer.

The processes disclosed herein can produce long-chain-branchedpolypropylene. Long-chain branching is achievable using the processesdisclosed herein regardless of whether additional α,ω-diene or otherdiene such as vinylnorbornene are used. In one form, less than 0.5 wt %diene is used in the combined feed to any polymerization train of thereactor bank. Alternatively, less than 0.4 wt %, or 0.3 wt %, or 0.2 wt%, or 1000 wt ppm, or 500 wt ppm, or 200 wt ppm, or 100 wt ppm are used.

In some forms, the processes disclosed herein involve using as acomonomer an α,ω-diene and the olefin/α,ω-diene copolymers resultingfrom that use. Additionally, the processes disclosed herein involve acopolymerization reaction of olefin monomers, wherein the reactionincludes propylene and ethylene copolymerization with an α,ω-diene andthe copolymers that are made. These copolymers may be employed in avariety of articles including, for example, films, fibers, such asspunbonded and melt blown fibers, fabrics, such as nonwoven fabrics, andmolded articles. More particularly, these articles include, for example,cast films, oriented films, injection molded articles, blow moldedarticles, foamed articles, foam laminates and thermoformed articles.

It should be noted that while linear α,ω-dienes are disclosed, otherdienes can also be employed to make polymers using the processesdisclosed herein. These would include branched, substituted α,ω-dienes,such as 2-methyl-1,9-decadiene; cyclic dienes, such as vinylnorbornene;or aromatic types, such as divinyl benzene.

Other forms include copolymers having from 98 to 99.999 weight % olefinunits, and from 0.001 to 2.000 weight % α,ω-diene units. Copolymer formsmay have a weight-average molecular weight from 50,000 to 2,000,000,crystallization temperatures from 50° C. to 140° C. and a melt flow rate(MFR) from 0.1 g/10 min to 1500 g/10 min. These forms display highcrystallization temperatures intrinsically, hence there is no need forexternally added nucleating agents.

In other forms, the copolymer includes from 90 to 99.999 weight % ofpropylene units, from 0.000 to 8 weight % of olefin units other thanpropylene units and from 0.001 to 2 weight % α,ω-diene units. Copolymerforms may have weight-average molecular weights from 20,000 to2,000,000, crystallization temperatures (without the addition ofexternal nucleating agents) from 115° C. to 135° C. and MFRs from 0.1g/10 min to 100 g/10 min. The accompanying olefin may be any of C₂-C₂₀α-olefins, diolefins (with one internal olefin) and their mixturesthereof. More specifically, olefins include ethylene, butene-1,pentene-1, hexene-1, heptene-1,4-methyl-1-pentene, 3-methyl-1-pentene,4-methyl-1-hexene, 5-methyl-1-hexene, 1-octene, 1-decene, 1-undecene,and 1-dodecene.

Copolymers of isotactic polypropylene made under supercriticalconditions include ethylene and C₄-C₁₂ comonomers such asbutene-1,3-methylpentene-1, hexene-1,4-methylpentene-1, octene-1, anddecene-1. The in-line blending processes disclosed herein can preparethese copolymers without the use of solvent or in an environment withlow solvent concentration.

Propylene polymers produced typically comprise 0 to 60 weight % of acomonomer, or 1 to 50 weight %, or 2 to 40 weight %, or 4 to 30 weight%, or 5 to 25 weight %, or 5 to 20 weight %, and have one or more of:

-   -   1. a heat of fusion, ΔH_(f), of 30 J/g or more, or 50 J/g or        more, or 60 or more, or 70 or more, or 80 or more, or 90 or        more, or 95 or more, or 100 or more, or 105 or more or an ΔH_(f)        of 30 J/g or less, or 20 J/g or 0;    -   2. a weight average molecular weight (as measured by GPC DRI) of        20,000 or more, or 30,000 to 1,000,000, or 50,000 to 500,000, or        50,000 to 400,000;    -   3. a melt flow rate of 0.1 g/10 min or more, or 0.5 g/10 min or        more, or 1.0 g/10 min or more, or between 0.1 and 10,000 g/10        min;    -   4. a melting peak temperature of 55° C. or more, or 75° C. or        more, or 100° C. or more, or 125° C. or more, or 150° C. or        more, between 145 and 165° C.;    -   5. an M_(w)/M_(n) (as measured by GPC DRI) of about 1.5 to 20,        or about 1.5 to 10, or 1.8 to 4.

In another form, the polymers produced by the processes disclosed hereinhave a melt viscosity of less than 10,000 centipoises at 180° C. asmeasured on a Brookfield viscometer, or between 1000 to 3000 cP for someforms (such as packaging and adhesives) or between 5000 and 10,000 cPfor other applications.

Polymer Blends and Polymer Additives:

Many different types of polymer blends may be made by the monomerrecycle processes for use with fluid phase in-line polymer blendingdisclosed herein. A major fraction of a blend is defined as 50% or moreby weight of the blend. A minor fraction of a blend is defined as lessthan 50% by weight of the blend. The monomer recycle processes for usewith fluid phase in-line polymer blending disclosed herein areadvantageous for producing olefin-based polymer blends in which thepolymers can be divided into two groups, P1 and P2, for which thefollowing conditions satisfied:

N(P1+P2)=N(P2) and N(P2)≧N(P1)

where N(P1+P2) is the number of monomers in the combined monomer pool ofP1 and P2 polymer groups, and N(P1) and N(P2) are the number of monomersin the first (P1) and second (P2) group of polymers, respectively.

In some forms, the polymer blends produced by the monomer recycleprocesses for use with fluid phase in-line polymer blending disclosedherein include two or more polymers, including but not limited to,thermoplastic polymer(s) and/or elastomer(s). In another advantageousform, the polymer blends produced by the monomer recycle processes foruse with fluid phase in-line polymer blending disclosed herein includetwo polymers (a homopolymer produced in one parallel reactor train and acopolymer produced in a second parallel reactor train) wherein thehomopolymer and copolymer have a common monomer and are thermoplasticpolymer(s) and/or elastomer(s).

A “thermoplastic polymer(s)” is a polymer that can be melted by heat andthen cooled with out appreciable change in properties. Thermoplasticpolymers typically include, but are not limited to, polyolefins,polyamides, polyesters, polycarbonates, polysulfones, polyacetals,polylactones, acrylonitrile-butadiene-styrene resins, polyphenyleneoxide, polyphenylene sulfide, styrene-acrylonitrile resins, styrenemaleic anhydride, polyimides, aromatic polyketones, or mixtures of twoor more of the above. Polyolefins include, but are not limited to,polymers comprising one or more linear, branched or cyclic C₂ to C₄₀olefins, polymers comprising propylene copolymerized with one or more C₂or C₄ to C₄₀ olefins, C₃ to C₂₀ alpha olefins, or C₃ to C₁₀ α-olefins.Also, polyolefins include, but are not limited to, polymers comprisingethylene including but not limited to ethylene copolymerized with a C₃to C₄₀ olefin, a C₃ to C₂₀ alpha olefin, propylene and/or butene.

“Elastomers” encompass all natural and synthetic rubbers, includingthose defined in ASTM D1566). Examples of useful elastomers include, butare not limited to, ethylene propylene rubber, ethylene propylene dienemonomer rubber, styrenic block copolymer rubbers (including SI, SIS, SB,SBS, SEBS and the like, where S=styrene, I=isobutylene, andB=butadiene), butyl rubber, halobutyl rubber, copolymers of isobutyleneand para-alkylstyrene, halogenated copolymers of isobutylene andpara-alkylstyrene, natural rubber, polyisoprene, copolymers of butadienewith acrylonitrile, polychloroprene, alkyl acrylate rubber, chlorinatedisoprene rubber, acrylonitrile chlorinated isoprene rubber,polybutadiene rubber (both cis and trans).

In another form, the polymer blends produced herein may include one ormore of isotactic polypropylene, highly isotactic polypropylene,syndiotactic polypropylene, atactic polypropylene, random copolymer ofpropylene and ethylene and/or butene and/or hexene, polybutene, ethylenevinyl acetate, low density polyethylene (density 0.915 to less than0.935 g/cm³), linear low density polyethylene, ultra low densitypolyethylene (density 0.86 to less than 0.90 g/cm³), very low densitypolyethylene (density 0.90 to less than 0.915 g/cm³), medium densitypolyethylene (density 0.935 to less than 0.945 g/cm³), high densitypolyethylene (density 0.945 to 0.98 g/cm³), ethylene vinyl acetate,ethylene methyl acrylate, copolymers of acrylic acid,polymethylmethacrylate or any other polymers polymerizable by ahigh-pressure free radical process, polyvinylchloride, polybutene-1,isotactic polybutene, ABS resins, ethylene-propylene rubber (EPR),vulcanized EPR, EPDM, block copolymer, styrenic block copolymers,polyamides, polycarbonates, PET resins, crosslinked polyethylene,polymers that are a hydrolysis product of EVA that equate to an ethylenevinyl alcohol copolymer, polymers of aromatic monomers such aspolystyrene, poly-1 esters, polyacetal, polyvinylidine fluoride,polyethylene glycols and/or polyisobutylene.

In another form, elastomers are blended using the processes disclosedherein to form rubber toughened compositions. In some forms, the rubbertoughened composition is a two (or more) phase system where theelastomer is a discontinuous phase and the polymer produced herein is acontinuous phase. This blend may be combined with tackifiers and/orother additives as described herein.

In another form, the polymer blends produced by the processes disclosedherein may include elastomers or other soft polymers to form impactcopolymers. In some forms, the blend is a two (or more) phase systemwhere the elastomer or soft polymer is a discontinuous phase and otherpolymer(s) is a continuous phase. The blends produced herein may becombined with tackifiers and/or other additives as described herein.

In some forms, the polymers blends disclosed herein include metallocenepolyethylenes (mPEs) or metallocene polypropylenes (mPPs). The mPE andmPP homopolymers or copolymers are typically produced using mono- orbis-cyclopentadienyl transition metal catalysts in combination with anactivator of aluminoxane and/or a non-coordinating anion in solution,slurry, high pressure or gas phase. The catalyst and activator may besupported or unsupported and the cyclopentadienyl rings may besubstituted or unsubstituted. Several commercial products produced withsuch catalyst/activator combinations are commercially available fromExxonMobil Chemical Company in Baytown, Tex. under the tradenamesEXCEED™, ACHIEVE™ and EXACT™. For more information on the methods andcatalysts/activators to produce such homopolymers and copolymers see WO94/26816; WO 94/03506; EPA 277,003; EPA 277,004; U.S. Pat. No.5,153,157; U.S. Pat. No. 5,198,401; U.S. Pat. No. 5,240,894; U.S. Pat.No. 5,017,714; CA 1,268,753; U.S. Pat. No. 5,324,800; EPA 129,368; U.S.Pat. No. 5,264,405; EPA 520,732; WO 92 00333; U.S. Pat. No. 5,096,867;U.S. Pat. No. 5,507,475; EPA 426 637; EPA 573 403; EPA 520 732; EPA 495375; EPA 500 944; EPA 570 982; WO91/09882; WO94/03506 and U.S. Pat. No.5,055,438.

In some forms the polymer blends produced by the processes disclosedherein include one polymer at from 10 to 99 weight %, based upon theweight of the polymers in the blend, or 20 to 95 weight %, or at least30 to 90 weight %, or at least 40 to 90 weight %, or at least 50 to 90weight %, or at least 60 to 90 weight %, or at least 70 to 90 weight %with one or more other polymers constituting the remainder of the blend.

In another form, in-line polymer blends are produced frompropylene-based polymers made at homogeneous polymerization conditions,particularly at bulk homogeneous polymerization conditions, such as bulkhomogeneous supercritical or bulk solution polymerization.

In another form, in-line polymer blends are produced frompropylene-based polymers made at homogeneous polymerization conditions,particularly at bulk homogeneous polymerization conditions, such as bulkhomogeneous supercritical or bulk solution polymerization, and comprisethe following:

-   (a) 10-20 wt % of isotactic polypropylene with 0.8-10,000 g/10 min    MFR and melting peak temperatures of 80-165° C. plus 80-90 wt %    crystallizable ethylene-propylene copolymer comprising 10-16 wt %    ethylene content and 0.8-100 g/10 min MFR or-   (b) 15-90 wt % of isotactic polypropylene with 0.8-10,000 g/10 min    MFR and melting peak temperatures of 80-165° C. plus 10-85 wt %    propylene copolymer of isotactic polypropylene crystallinity    comprising 1-20 wt % ethylene or 1-40 wt % hexene-1 or 1-30 wt    butene-1 content and 0.8-100 g/10 min MFR or-   (c) 10-30 wt % of isotactic polypropylene with 0.8-10,000 g/10 min    MFR and melting peak temperatures of 80-165° C. plus 90-70 wt %    low-crystallinity (0-30 J/g) homo- or copolymer with MFR of 0.8-500    g/10 min.

The in-line polymer blends produced by the process disclosed herein maybe also blended with other polymers and additives using the in-lineblending process for other polymers and additives depicted in FIGS. 10and 11, in an extrusion process downstream of in-linepolymerization/separation/blending processes disclosed herein, orblended in an off-line compounding process.

Any of the above polymers included in the in-line polymer blendsproduced by the processes disclosed herein may be functionalized.Functionalized means that the polymer has been contacted with anunsaturated acid or anhydride. Forms of unsaturated acids or anhydridesinclude any unsaturated organic compound containing at least one doublebond and at least one carbonyl group. Representative acids includecarboxylic acids, anhydrides, esters and their salts, both metallic andnon-metallic. The organic compound contains an ethylenic unsaturationconjugated with a carbonyl group (—C═O). Non-limiting examples includemaleic, fumaric, acrylic, methacrylic, itaconic, crotonic, alpha-methylcrotonic, and cinnamic acids as well as their anhydrides, esters andsalt derivatives. Maleic anhydride is one particular form. Theunsaturated acid or anhydride is present at about 0.1 weight % to about5 weight %, or at about 0.5 weight % to about 4 weight %, or at about 1to about 3 weight %, based upon the weight of the hydrocarbon resin andthe unsaturated acid or anhydride.

Tackifiers may also be blended either in-line by the processes disclosedherein (see FIGS. 10 and 11), in-line via an extrusion processdownstream of in-line polymerization/separation/blending processesdisclosed herein, or in an off-line compounding process. Examples ofuseful tackifiers include, but are not limited to, aliphatic hydrocarbonresins, aromatic modified aliphatic hydrocarbon resins, hydrogenatedpolycyclopentadiene resins, polycyclopentadiene resins, gum rosins, gumrosin esters, wood rosins, wood rosin esters, tall oil rosins, tall oilrosin esters, polyterpenes, aromatic modified polyterpenes, terpenephenolics, aromatic modified hydrogenated polycyclopentadiene resins,hydrogenated aliphatic resin, hydrogenated aliphatic aromatic resins,hydrogenated terpenes and modified terpenes, and hydrogenated rosinesters. In some embodiments the tackifier is hydrogenated. In otherembodiments the tackifier is non-polar. Non-polar tackifiers aresubstantially free of monomers having polar groups. The polar groups aregenerally not present; however, if present, they are not present at morethat 5 weight %, or not more that 2 weight %, or no more than 0.5 weight%. In some embodiments, the tackifier has a softening point (Ring andBall, as measured by ASTM E-28) of 80° C. to 140° C., or 100° C. to 130°C. In some embodiments the tackifier is functionalized. Byfunctionalized is meant that the hydrocarbon resin has been contactedwith an unsaturated acid or anhydride. Useful unsaturated acids oranhydrides include any unsaturated organic compound containing at leastone double bond and at least one carbonyl group. Representative acidsinclude carboxylic acids, anhydrides, esters and their salts, bothmetallic and non-metallic. The organic compound may contain an ethylenicunsaturation conjugated with a carbonyl group (—C═O). Non-limitingexamples include maleic, fumaric, acrylic, methacrylic, itaconic,crotonic, alpha-methyl crotonic, and cinnamic acids as well as theiranhydrides, esters and salt derivatives. Maleic anhydride isparticularly useful. The unsaturated acid or anhydride may be present inthe tackifier at about 0.1 weight % to 10 weight %, or at 0.5 weight %to 7 weight %, or at 1 to 4 weight %, based upon the weight of thehydrocarbon resin and the unsaturated acid or anhydride.

The tackifier, if present, is typically present at 1 weight % to 50weight %, based upon the weight of the blend, or 10 weight % to 40weight %, or 20 weight % to 40 weight %. Generally however, tackifier isnot present, or if present, is present at less than 10 weight %, or lessthan 5 weight %, or at less than 1 weight %.

In another form, the polymer blends produced by the processes disclosedherein further comprise a crosslinking agent. The crosslinking agent maybe blended either in-line by the processes disclosed herein (see FIGS.10 and 11), in-line via an extrusion process downstream of in-linepolymerization/separation/blending processes disclosed herein, or in anoff-line compounding process. Useful crosslinking agents include thosehaving functional groups that can react with the acid or anhydride groupand include alcohols, multiols, amines, diamines and/or triamines.Non-limiting examples of crosslinking agents useful include polyaminessuch as ethylenediamine, diethylenetriamine, hexamethylenediamine,diethylaminopropylamine, and/or menthanediamine.

In another form, the polymer blends produced by the processes disclosedherein, and/or blends thereof, further comprise typical additives knownin the art such as fillers, cavitating agents, antioxidants,surfactants, adjuvants, plasticizers, block, antiblock, colormasterbatches, pigments, dyes, processing aids, UV stabilizers,neutralizers, lubricants, waxes, nucleating agents and/or clarifyingagents. These additives may be present in the typically effectiveamounts well known in the art, such as 0.001 weight % to 10 weight %.These additive may be blended either in-line by the processes disclosedherein (see FIGS. 10 and 11), in-line via an extrusion processdownstream of in-line polymerization/separation/blending processesdisclosed herein, or in an off-line compounding process.

Useful fillers, cavitating agents and/or nucleating agents includetitanium dioxide, calcium carbonate, barium sulfate, silica, silicondioxide, carbon black, sand, glass beads, mineral aggregates, talc, clayand the like. Nucleating agents of the non-clarifying type include, butare not limited to, sodium benzoate, Amfine NA 11, Amfine NA 21, andMilliken HPN 68.

Useful antioxidants and UV stablilizers include phenolic antioxidants,such as Irganox 1010, Irganox 1076 both available from Ciba-Geigy. Oilsmay include paraffinic or naphthenic oils such as Primol 352, or Primol876 available from ExxonMobil Chemical France, S.A. in Paris, France.The oils may include aliphatic naphthenic oils, white oils or the like.

Plasticizers and/or adjuvants may include mineral oils, polybutenes,phthalates and the like. The plasticizers may include phthalates such asdiisoundecyl phthalate (DIUP), diisononylphthalate (DINP),dioctylphthalates (DOP) and polybutenes, such as Parapol 950 and Parapol1300 available from ExxonMobil Chemical Company in Houston Tex.Additional plasticizers include those disclosed in WO0118109A1, U.S.patent application Ser. No. 10/640,435, and U.S. patent application Ser.No. 11/177,004, which are incorporated by reference herein with regardto plasticizer compositions and blending thereof.

Useful processing aids, lubricants, waxes, and/or oils include lowmolecular weight products such as wax, oil or low M_(n) polymer, (lowmeaning below M_(n) of 5000, or below 4000, or below 3000, or below2500). Useful waxes include polar or non-polar waxes, functionalizedwaxes, polypropylene waxes, polyethylene waxes, and wax modifiers.

Useful functionalized waxes include those modified with an alcohol, anacid, or a ketone. Functionalized means that the polymer has beencontacted with an unsaturated acid or anhydride. Useful unsaturatedacids or anhydrides include any unsaturated organic compound containingat least one double bond and at least one carbonyl group. Representativeacids include carboxylic acids, anhydrides, esters and their salts, bothmetallic and non-metallic. The organic compound may contain an ethylenicunsaturation conjugated with a carbonyl group (—C═O). Non-limitingexamples include maleic, fumaric, acrylic, methacrylic, itaconic,crotonic, alpha-methyl crotonic, and cinnamic acids as well as theiranhydrides, esters and salt derivatives. Maleic anhydride isparticularly useful. The unsaturated acid or anhydride may be present at0.1 weight % to 10 weight %, or at 0.5 weight % to 7 weight %, or at 1to 4 weight %, based upon the weight of the hydrocarbon resin and theunsaturated acid or anhydride. Examples include waxes modified by methylketone, maleic anhydride or maleic acid. Low Mn polymers includepolymers of lower alpha olefins such as propylene, butene, pentene,hexene and the like. A useful polymer includes polybutene having an Mnof less than 1000 g/mol. An example of such a polymer is available underthe trade name PARAPOL™ 950 from ExxonMobil Chemical Company. PARAPOL™950 is a liquid polybutene polymer having an Mn of 950 g/mol and akinematic viscosity of 220 cSt at 100° C., as measured by ASTM D 445.

Useful clarifying agents include, but are not limited to, thebenzalsorbitol family of clarifiers, and more particularlydibenzalsorbitol (Millad 3905), di-p-methylbenzalsorbitol (Milliad3940), and bis-3,4-dimethylbenzalsorbitol (Milliad 3988).

Applications:

The polymer blends produced by the processes disclosed herein aretypically used in any known thermoplastic or elastomer application.Non-limiting examples include uses in molded parts, films, tapes,sheets, tubing, hose, sheeting, wire and cable coating, adhesives, shoesoles, bumpers, gaskets, bellows, films, fibers, elastic fibers,nonwovens, spunbonds, sealants, surgical gowns and medical devices.

Applicants have attempted to disclose all embodiments and applicationsof the disclosed subject matter that could be reasonably foreseen.However, there may be unforeseeable, insubstantial modifications thatremain as equivalents. While the present invention has been described inconjunction with specific, exemplary embodiments thereof, it is evidentthat many alterations, modifications, and variations will be apparent tothose skilled in the art in light of the foregoing description withoutdeparting from the spirit or scope of the present disclosure.Accordingly, the present disclosure is intended to embrace all suchalterations, modifications, and variations of the above detaileddescription.

All patents, test procedures, and other documents cited herein,including priority documents, are fully incorporated by reference to theextent such disclosure is not inconsistent with this invention and forall jurisdictions in which such incorporation is permitted.

When numerical lower limits and numerical upper limits are listedherein, ranges from any lower limit to any upper limit are contemplated.All numerical values within the detailed description and the claimsherein are also understood as modified by “about.”

1. An in-line blending process for polymers comprising: (a) providingtwo or more reactor trains configured in parallel and two or morehigh-pressure separators downstream fluidly connected to the two or morereactor trains configured in parallel; (b) contacting in the two or morereactor trains configured in parallel 1) olefin monomers having two ormore carbon atoms, 2) one or more catalyst systems, 3) optional one ormore comonomers, 4) optional one or more scavengers, and 5) optional oneor more diluents or solvents, wherein at least one of the reactor trainsconfigured in parallel is at a temperature above the solid-fluidphase-transition temperature of the polymerization system and a pressureno lower than 10 MPa below the cloud point pressure of thepolymerization system and less than 1500 MPa, wherein the polymerizationsystem for each reactor train is in its dense fluid state and comprisesthe olefin monomers, any comonomer present, any diluent or solventpresent, any scavenger present, and the polymer product, and wherein thecatalyst system for each reactor train comprises one or more catalystprecursors, one or more activators, and optionally, one or more catalystsupports; (c) forming an unreduced reactor effluent including ahomogenous fluid phase polymer-monomer mixture in each reactor trainconfigured in parallel, wherein one or more of the parallel reactortrains define a first group (G1) of reactor trains and another one ormore of the parallel reactor trains define a second group (G2) ofreactor trains, wherein the number of monomers (N) in the feed monomerpools for G1 (N(G1)), G2 (N(G2)), and for the combined feed monomer poolof G1 and G2 (N(G1+G2)) are related as follows: N(G1+G2)=N(G2) andN(G1)≦N(G2), and wherein the feed monomer pool in each reactor train ofG2 are the same; (d) passing the unreduced reactor effluents from one ormore of the G1 reactor trains through one or more high-pressureseparators, maintaining the temperature and pressure within the one ormore high-pressure separators above the solid-fluid phase transitionpoint but below the cloud point pressure and temperature to formfluid-fluid two-phase systems in each of the one or more high-pressureseparators with each two-phase system comprising a polymer-enrichedphase and a monomer-rich phase, and separating the monomer-rich phasefrom the polymer-enriched phase in each of the one or more high-pressureseparators to form one or more separated monomer-rich phases and one ormore polymer-enriched phases; (e) passing the one or morepolymer-enriched phases from the one or more high-pressure separators of(d), any unreduced reactor effluents of G1 not passing through the oneor more high-pressure separators of G1, and the unreduced reactoreffluents of the G2 reactor trains through another high-pressureseparator for product blending and product-feed separation; (f)maintaining the temperature and pressure within the anotherhigh-pressure separator of (e) above the solid-fluid phase transitionpoint but below the cloud point pressure and temperature to form afluid-fluid two-phase system comprising a polymer-rich blend phase and amonomer-rich phase; (g) separating the monomer-rich phase from thepolymer-rich blend phase in the another high-pressure separator of (e)to form a polymer-rich blend and a separated monomer-rich phase; (h)recycling the one or more separated monomer-rich phases of (d) to one ormore reactor trains of G1; and (i) recycling the separated monomer-richphase of (g) to one or more reactor trains of G2.
 2. The process ofclaim 1 wherein the combined flow rates of each monomer in the one ormore polymer-enriched phases from (d) plus the combined flow rate ofeach monomer in any unreduced reactor effluents of G1 not passingthrough the one or more high-pressure separators of G1 minus thecombined flow rate of each monomer purged from the process of claim 1 isless than or equal to the combined conversion rate of the correspondingmonomer in the G2 reactor trains.
 3. The process of claim 1 wherein G1includes one reactor train and G2 includes one reactor train.
 4. Theprocess of claim 1 wherein at least one of the reactor trains of each ofG1 and G2 includes propylene.
 5. The process of claim 1 wherein theoptional one or more comonomers of (b) comprise one or more of ethylene,propylene, butenes, hexenes, octenes, decenes, or dodecenes.
 6. Theprocess of claim 1 wherein the one or more reactor trains of G1polymerize a homopolymer and the one or more reactor trains of G2polymerize a copolymer.
 7. The process of claim 6 wherein thehomopolymer is chosen from isotactic polypropylene, syndiotacticpolypropylene, atactic polypropylene, polyethylene, polybutene,polyhexene, polyoctene, polydecene, and polystyrene.
 8. The process ofclaim 6 wherein the copolymer is chosen from ethylene-propylene,propylene-butene-1, propylene-hexene-1, propylene-octene-1,propylene-decene-1, ethylene-butene-1, ethylene-hexene-1,ethylene-octene-1, ethylene-propylene-butene-1,ethylene-propylene-hexene-1, ethylene-propylene-octene-1, andethylene-propylene-decene-1 copolymers.
 9. The process of claim 1further comprising removing low molecular weight oligomers, lowmolecular weight polymers, solvents/diluents or combinations thereoffrom the separated monomer-rich phases of (d) and/or (g).
 10. Theprocess of claim 9 wherein the low-molecular weight oligomers, lowmolecular weight polymers, solvents/diluents or combinations thereof areremoved through the use of at least one knock-out pot, at least oneseparation tower, or a combination thereof.
 11. The process of claim 1further comprising feeding the combined one or more polymer-enrichedphases from the one or more high-pressure separators of (d), theunreduced reactor effluents of G1 not passing through the one or morehigh-pressure separators of G1, and the unreduced reactor effluents ofthe G2 reactor trains through one or more static mixers.
 12. The processof claim 1 further comprising providing one or more storage tanks, andfeeding from the one or more storage tanks one or more polymers and/orpolymer additives to the process after (c).
 13. The process of claim 1further comprising one or more pressure letdown valves for each reactortrain positioned before the two or more high-pressure separators of (a),and one or more pressure letdown valves between the two or morehigh-pressure separators of (a).
 14. The process of claim 1 furthercomprising feeding the polymer-rich blend of (g) to one or morelow-pressure separators serially connected to further separate themonomers and other volatiles to form a further-enriched polymer blendand one or more monomer-rich streams.
 15. The process of 14 furthercomprising feeding the further-enriched polymer blend to a coupleddevolatizer to further separate other volatiles to form a polymer blend.16. The process of claim 15 wherein the coupled devolatilizer operatesunder vacuum enabling the further-enriched polymer blend to flash offthe monomers and other volatiles.
 17. The process of claim 15 whereinthe coupled devolatilizer is a devolatizing extruder.
 18. The process ofclaim 17 wherein one or more polymer additives are added to the polymerblend at one or more high-pressure separators, one or more low-pressureseparators, the devolatizing extruder or combinations thereof.
 19. Theprocess of claim 1 wherein the two or more high-pressure separators of(a) are gravimetric separation vessels.
 20. The process of claim 13wherein the pressure is dropped at a rate of at least about 6 MPa/sec.21. The process of claims 12 or 18 wherein the one or more polymeradditives are chosen from waxes, polyalfaolefins, antioxidants,plasticizers, nucleating and clarifying agents, slip agents, flameretardants, heat and UV stabilizers, antiblocking agents, fillers,reinforcing fibers, antistatic agents, lubricating agents, coloringagents, foaming agents, and combinations thereof.
 22. The process ofclaim 1 wherein one or more of the reactor trains configured in parallelof (a) comprise a tubular reactor, a stirred autoclave reactor, a loopreactor, or combinations thereof.
 23. The process of claim 1 wherein oneor more of the reactor trains configured in parallel of (a) comprise twoor more reactors in series wherein the two or more reactors in seriescomprise a tubular reactor followed by a stirred autoclave reactor or atubular reactor followed by a loop reactor.
 24. An in-line blendingprocess for polymers comprising: (a) providing two or more reactortrains configured in parallel and a high-pressure separator downstreamfluidly connected to the two or more reactor trains configured inparallel; (b) contacting in the two or more reactor trains configured inparallel 1) olefin monomers having two or more carbon atoms, 2) one ormore catalyst systems, 3) optional one or more comonomers, 4) optionalone or more scavengers, and 5) optional one or more diluents orsolvents, wherein at least one of the reactor trains configured inparallel is at a temperature above the solid-fluid phase-transitiontemperature of the polymerization system and a pressure no lower than 10MPa below the cloud point pressure of the polymerization system and lessthan 1500 MPa, wherein the polymerization system for each reactor trainis in its dense fluid state and comprises the olefin monomers, anycomonomer present, any diluent or solvent present, any scavengerpresent, and the polymer product, and wherein the catalyst system foreach reactor train comprises one or more catalyst precursors, one ormore activators, and optionally, one or more catalyst supports; (c)forming a reactor effluent including a homogeneous fluid phasepolymer-monomer mixture in each reactor train configured in parallel,wherein one or more of the parallel reactor trains define a first group(G1) of reactor trains and another one or more of the parallel reactortrains define a second group (G2) of reactor trains, wherein the numberof monomers (N) in the feed monomer pools for G1 (N(G1)), G2 (N(G2)),and for the combined feed monomer pool of G1 and G2 (N(G1+G2)) arerelated as follows: N(G1+G2)=N(G2) and N(G1)≦N(G2), and wherein the feedmonomer pool in each reactor train of G2 are the same; (d) passing thereactor effluent comprising the homogeneous fluid phase polymer-monomermixture from each reactor train of G1 and G2 through the high-pressureseparator for product blending and product-feed separation; (e)maintaining the temperature and pressure within the high-pressureseparator above the solid-fluid phase transition point but below thecloud point pressure and temperature to form a fluid-fluid two-phasesystem comprising a polymer-rich blend phase and a monomer-rich phase;(f) separating the monomer-rich phase from the polymer-rich blend phasein the high-pressure separator to form a polymer-rich blend and aseparated monomer-rich phase; and (g) recycling the separatedmonomer-rich phase of (f) to one or more reactor trains of G2.
 25. Theprocess of claim 24 wherein the combined flow rates of each monomer inthe reactor effluent from each reactor train of G1 minus the combinedflow rates of each monomer purged from the process of claim 24 is lessthan or equal to the combined conversion rate of the correspondingmonomer in the G2 reactor trains.
 26. The process of claim 24 wherein G1includes one reactor train and G2 includes one reactor train.
 27. Theprocess of claim 24 wherein at least one of the reactor trains of eachof G1 and G2 includes propylene.
 28. The process of claim 24 wherein theoptional one or more comonomers of (b) comprise one or more of ethylene,propylene, butenes, hexenes, octenes, decenes, or dodecenes.
 29. Theprocess of claim 24 wherein the one or more reactor trains of G1polymerize a homopolymer and the one or more reactor trains of G2polymerize a copolymer.
 30. The process of claim 29 wherein thehomopolymer is chosen from isotactic polypropylene, syndiotacticpolypropylene, atactic polypropylene, polyethylene, polybutene,polyhexene, polyoctene, polydecene, and polystyrene.
 31. The process ofclaim 29 wherein the copolymer is chosen from ethylene-propylene,propylene-butene-1, propylene-hexene-1, propylene-octene-1,propylene-decene-1, ethylene-butene-1, ethylene-hexene-1,ethylene-octene-1, ethylene-propylene-butene-1,ethylene-propylene-hexene-1, ethylene-propylene-octene-1 copolymers, andethylene-propylene-decene-1.
 32. The process of claim 24 furthercomprising removing low molecular weight oligomers, low molecular weightpolymers, solvents/diluents or combinations thereof from the separatedmonomer-rich phase of (f).
 33. The process of claim 32 wherein thelow-molecular weight oligomers, low molecular weight polymers,solvents/diluents or combinations thereof are removed through the use ofat least one knock-out pot, at least one separation tower, or acombination thereof.
 34. The process of claim 24 further comprisingfeeding the combined reactor effluents of G1 and/or G2 through one ormore static mixers.
 35. The process of claim 24 further comprisingproviding one or more storage tanks, and feeding from the one or morestorage tanks one or more polymers and/or polymer additives to theprocess after (c).
 36. The process of claim 24 further comprising one ormore pressure letdown valves for each reactor train positioned beforethe high-pressure separator.
 37. The process of claim 24 furthercomprising feeding the polymer-rich blend of (f) to one or morelow-pressure separators serially connected to further separate themonomers and other volatiles to form a further-enriched polymer blendand one or more monomer-rich streams.
 38. The process of 37 furthercomprising feeding the further-enriched polymer-rich blend to a coupleddevolatizer to further separate other volatiles to form a polymer blend.39. The process of claim 38 wherein the coupled devolatilizer operatesunder vacuum enabling the further-enriched polymer-rich blend to flashoff the monomers and other volatiles.
 40. The process of claim 38wherein the coupled devolatilizer is a devolatizing extruder.
 41. Theprocess of claim 38 wherein one or more polymer additives are added tothe polymer blend at the high-pressure separator, one or morelow-pressure separators, the devolatizing extruder or combinationsthereof.
 42. The process of claim 24 wherein the high-pressure separatorof (d) is a gravimetric separation vessel.
 43. The process of claim 36wherein the pressure is dropped at a rate of at least about 6 MPa/sec.44. The process of claims 35 or 41 wherein the one or more polymeradditives are chosen from waxes, polyalfaolefins, antioxidants,plasticizers, nucleating and clarifying agents, slip agents, flameretardants, heat and UV stabilizers, antiblocking agents, fillers,reinforcing fibers, antistatic agents, lubricating agents, coloringagents, foaming agents, and combinations thereof.
 45. The process ofclaim 24 wherein one or more of the reactor trains configured inparallel of (a) comprise a tubular reactor, a stirred autoclave reactor,a loop reactor, or combinations thereof.
 46. The process of claim 24wherein one or more of the reactor trains configured in parallel of (a)comprise two or more reactors in series wherein the two or more reactorsin series comprise a tubular reactor followed by a stirred autoclavereactor or a tubular reactor followed by a loop reactor.
 47. A in-lineblending process for polymers comprising: (a) providing a first group(G1) of one or more reactor trains and a second group (G2) of one ormore reactor trains, wherein G1 polymerizes a polypropylene homopolymerand G2 polymerizes a polypropylene copolymer, wherein one high-pressureseparator is fluidly connected to G1 and another high-pressure separatoris fluidly connected to G2, and wherein the G1 reactor trains areconfigured parallel to and fluidly connected to the G2 reactor trains;(b) contacting in each of the reactor trains of G1 and G2: 1) propylenemonomer, 2) one or more catalyst systems, 3) one or more comonomers ineach of the reactor trains of G2, 4) optional one or more scavengers,and 5) optional one or more diluents or solvents, wherein each of thereactor trains for G1 and G2 are at a temperature above the solid-fluidphase-transition temperature of the polymerization system and a pressureno lower than 10 MPa below the cloud point pressure of thepolymerization system and less than 1500 MPa and have at least onecommon monomer, wherein the polymerization system for each reactor trainof G1 and G2 is in its dense fluid state and comprises propylenemonomer, any diluent or solvent present, any scavenger present, and thepolymer product, and the polymerization system for the second reactortrain further comprises one or more comonomers, and wherein the catalystsystem for each reactor train of G1 and G2 comprises one or morecatalyst precursors, one or more activators, and optionally, one or morecatalyst supports; (c) forming an unreduced reactor effluent including ahomogenous fluid phase polymer-monomer mixture in each reactor train ofG1 and G2; (d) passing the unreduced reactor effluents from the one ormore of the G1 reactor trains through a high-pressure separator whilemaintaining the temperature and pressure within the high-pressureseparator above the solid-fluid phase transition point but below thecloud point pressure and temperature to form a fluid-fluid two-phasesystems comprising a polymer-enriched phase and a monomer-rich phase,and separating the monomer-rich phase from the polymer-enriched phase inthe high-pressure separator to form a separated monomer-rich phase and apolymer-enriched phase; (e) passing the polymer-enriched phase from thehigh-pressure separator of (d) and the unreduced reactor effluents ofthe G2 reactor trains through another high-pressure separator forproduct blending and product-feed separation; (f) maintaining thetemperature and pressure within the another high-pressure separator of(e) above the solid-fluid phase transition point but below the cloudpoint pressure and temperature to form a fluid-fluid two-phase systemcomprising a polymer-rich blend phase and a monomer-rich phase; (g)separating the monomer-rich phase from the polymer-rich blend phase inthe another high-pressure separator of (e) to form a polymer-rich blendand a separated monomer-rich phase; (h) recycling the separatedmonomer-rich phase of (d) to the one or more reactor trains of G1; and(i) recycling the separated monomer-rich phase of (g) to one or morereactor trains of G2.
 48. The process of claim 47 wherein thepolypropylene copolymer is chosen from ethylene-propylene,propylene-butene-1, propylene-hexene-1, propylene-octene-1,propylene-decene-1, ethylene-propylene-butene-1,ethylene-propylene-hexene-1, ethylene-propylene-octene-1, andethylene-propylene-decene-1 copolymers.
 49. The process of claim 47wherein the combined flow rate of propylene monomer in thepolymer-enriched phase from (d) minus the combined flow rate ofpropylene monomer purged from the process of claim 47 is less than orequal to the combined conversion rate of the propylene monomer in the G2reactor trains of (a).
 50. A in-line blending process for polymerscomprising: (a) providing a first group (G1) of one or more reactortrains and a second group (G2) of one or more reactor trains, wherein G1polymerizes a polypropylene homopolymer and G2 polymerizes apolypropylene copolymer, wherein a high-pressure separator is fluidlyconnected to G1 and G2, and wherein the G1 reactor trains are configuredparallel to and fluidly connected to the G2 reactor trains; (b)contacting in each of the reactor trains of G1 and G2:1) propylenemonomer, 2) one or more catalyst systems, 3) one or more comonomers ineach of the reactor trains of G2, 4) optional one or more scavengers,and 5) optional one or more diluents or solvents, wherein each of thereactor trains for G1 and G2 are at a temperature above the solid-fluidphase-transition temperature of the polymerization system and a pressureno lower than 10 MPa below the cloud point pressure of thepolymerization system and less than 1500 MPa and have at least onecommon monomer, wherein the polymerization system for each reactor trainof G1 and G2 is in its dense fluid state and comprises propylenemonomer, any diluent or solvent present, any scavenger present, and thepolymer product, and the polymerization system for the second reactortrain further comprises one or more comonomers, and wherein the catalystsystem for each reactor train of G1 and G2 comprises one or morecatalyst precursors, one or more activators, and optionally, one or morecatalyst supports; (c) forming a reactor effluent including ahomogeneous fluid phase polymer-monomer mixture in each reactor train ofG1 and G2; (d) passing the reactor effluents from each reactor train ofG1 and G2 through the high-pressure separator for product blending andproduct-feed separation; (e) maintaining the temperature and pressurewithin the high-pressure separator above the solid-fluid phasetransition point but below the cloud point pressure and temperature toform a fluid-fluid two-phase system comprising a polymer-rich blendphase and a monomer-rich phase; (f) separating the monomer-rich phasefrom the polymer-rich blend phase in the high-pressure separator to forma polymer-rich blend and a separated monomer-rich phase; and (g)recycling the separated monomer-rich phase of (f) to the one or morereactor trains of G2.
 51. The process of claim 50 wherein thepolypropylene copolymer is chosen from ethylene-propylene,propylene-butene-1, propylene-hexene-1, propylene-octene-1,propylene-decene-1, ethylene-propylene-butene-1,ethylene-propylene-hexene-1, ethylene-propylene-octene-1, andethylene-propylene-decene-1 copolymers.
 52. The process of claim 50wherein the combined flow rate of propylene monomer in the reactoreffluent from all reactor trains of G1 minus the combined flow rate ofpropylene monomer purged from the process of claim 50 is less than orequal to the combined conversion rate of propylene monomer in the G2reactor trains of (a).